Sajtar, Eric T., Toolbox to Evaluate Treatment Technologies for PRB CBM Water M.S.,

Department of Civil and Architectural Engineering, May, 2008.

Coalbed methane (CBM) co-produced water in the Powder River Basin (PRB) can contain,

among other constituents, sodium ranging from 110-800 mg/L, total dissolved solids ranging

from 270-2,010 mg/L, and sodium adsorption ratios ranging from 5.7-33. Due to excess sodium

concentrations, untreated CBM waters may negatively affect aquatic ecosystems, damage soil

characteristics, or contaminate aquifers. These issues can make it necessary to treat these waters

before discharge or use. The scope of this work has been to develop the Coalbed Methane

Demineralization Treatment Cost (CBM-DTC) Toolbox that provides initial demineralization

technology selection for removing sodium and other constituents from PRB CBM waters based

on influent water properties and effluent requirements. The toolbox has been constructed by

developing performance and cost functions for selected demineralization technologies (ion

exchange, reverse osmosis, electrodialysis reversal, and high efficiency electro-pressure

membranes). Sensitivity analyses were performed using the toolbox to determine which

parameters have the largest impact on treatment cost. The amount of sodium removed was found

to be the most important parameter followed by influent flow rate, facility life and brine disposal

cost. Further analyses have shown treatment costs to range from $0.036 per barrel (bbl) of

influent water to $0.190/bbl depending on technology and conditions selected.

TOOLBOX TO EVALUATE TREATMENT TECHNOLOGIES

FOR PRB CBM WATER

by

Eric T. Sajtar

A thesis submitted to the Department of Civil and Architectural Engineering

and The Graduate School of The University of Wyoming

in partial fulfillment of the requirements

for the degree of

MASTER OF SCIENCE

in

ENVIRONMENTAL ENGINEERING

Laramie, Wyoming

May 2008
ii

ACKNOWLEDGMENTS

I would like to thank the United States Department of Energy National Energy Technology

Laboratory (Award No. DE-FC26-06NT15568) for funding this work and my advisor Dr. David

Bagley for discussing and reviewing my work throughout the course of this project. I would also

like to thank my girl for her continued support. Finally, I would like to thank myself for sticking

it out and completing this work.
iii

TABLE OF CONTENTS Page

Acknowledgements ii

Table of Contents iii

List of Figures ix

List of Tables xiii

1 INTRODUCTION 1

1.1 Background 1

1.2 Objectives 3

2 WATER CHARACTERIZATION 5

2.1 Produced Water Characteristics 5

2.2 Water Use Requirements 7

2.2.1 Surface Discharge 7

2.2.2 Other Uses 8

3 TREATMENT TECHNOLOGIES 12

3.1 Overview 12

3.1.1 Process 12

3.1.2 Costs 15

3.2 Pretreatment 16

3.2.1 Methane Removal 16

3.2.2 Equalization Storage 17

3.2.3 Iron Removal 18

3.2.3.1 Iron Pretreatment Capital Costs 20

3.2.3.2 Iron Pretreatment Operating and Maintenance Costs 21
iv

3.2.4 Acid Addition 22

3.3 Reverse Osmosis 24

3.3.1 Overview 24

3.3.2 Water Recovery 28

3.3.3 Capital Costs 31

3.3.4 Operating and Maintenance 32

3.3.5 Energy Recovery and Savings 36

3.4 Ion Exchange 37

3.4.1 Overview 37

3.4.2 Process 39

3.4.3 Design Requirements 41

3.4.3.1 Regeneration 42

3.4.3.2 Waste Stream 43

3.4.4 Configurations 44

3.4.4.1 Regeneration Configurations 44

3.4.4.2 Continuous bed vs. Non-continuous bed 45

3.4.4.3 Water-resin contacting methods 47

3.4.4.4 Waste Stream Management 48

3.4.5 PRB Ion Exchange Vendors 51

3.4.6 Costs 52

3.4.6.1 Capital Costs 53

3.4.6.2 Operating and Maintenance Costs 54

3.5 Electrodialysis 57
v

3.5.1 Overview 57

3.5.2 Process 58

3.5.3 Electrodialysis Reversal (EDR) 60

3.5.4 Chemical Rejection 60

3.5.5 Water Recovery 61

3.5.6 Capital Costs 62

3.5.7 Operating and Maintenance Costs 64

3.6 HEEPM 67

3.6.1 Process 67

3.6.2 HEEPM Costs 68

3.6.2.1 Capital Costs 69

3.6.2.2 Operating and Maintenance Costs 70

4 BRINE MANAGEMENT 74

4.1 Overview 74

4.2 Deep Well Injection 74

4.2.1 Overview 74

4.2.2 Costs 75

4.3 Evaporation Ponds 77

4.3.1 Overview 77

4.3.2 Parameters 77

4.3.3 Costs 79

4.4 Evaporation Crystallization 80

4.4.1 Overview 80
vi

4.4.2 Process 80

4.4.3 Costs 83

4.4.3.1 Evaporator 84

4.4.3.2 Crystallizer 87

4.5 Other Options 89

4.5.1 Bipolar Electrodialysis 89

4.5.2 Membrane Distillation 91

5 TOOLBOX 93

5.1 Overview 93

5.2 Influent and Effluent 94

5.3 Technologies 96

5.3.1 Input 96

5.3.2 Output 97

5.4 Brine 98

5.4.1 Input 98

5.4.2 Output 99

6 SENSITIVITY ANALYSIS 101

6.1 Overview 101

6.2 Sensitivity Analysis Setup 101

6.3 Results and Discussion 103

6.3.1 Background 103

6.3.2 Reverse Osmosis 105

6.3.3 Ion Exchange 106
vii

6.3.4 Electrodialysis Reversal 108

6.3.5 HEEPM 109

6.4 Summary 111

7. TECHNOLOGY COST EFFECTIVENESS AND SELECTION 112

7.1 Overview 112

7.2 Scenarios 112

7.3 Results and Discussion 114

7.3.1 Scenarios 1 and 2: Facility Life 114

7.3.2 Scenario 3: Influent Flow Rate 118

7.3.3 Scenario 4: Brine Disposal Cost 121

7.3.4 Scenario 5: Electrical Cost 124

7.3.5 Scenarios 6 and 7: Acid Cost 126

7.4 Summary 130

8. SUMMARY, CONCLUSIONS AND RECOMMENDATIONS 132

8.1 Summary 132

8.2 Conclusions 133

8.3 Recommendations 134

9. REFERENCES 136

10. APPENDICES 145

Appendix A: Treatment Technology Data and Analysis 145

Appendix A-P: Pretreatment 145

Appendix A-RO: Reverse Osmosis 147

Appendix A-IX: Ion Exchange 151
viii

Appendix A-ED: Electrodialysis 152

Appendix A-H: HEEPM 155

Appendix B: Brine Management Analysis 161

Appendix C: List of Reviewed WYPDES Permits 162

Appendix D: Factorial design experimental setups and toolbox results 163

Appendix E: Units and Unit Conversions 170
ix

LIST OF FIGURES

Figure Page

Figure 3-1: General treatment flow schematic 13

Figure 3-2: Schematic of degasser for methane removal 17

Figure 3-3: Iron removal capital vs. flow rate 21

Figure 3-4: Reverse osmosis treatment process

A) Osmosis B) Reverse osmosis 25

Figure 3-5: Reverse osmosis water recovery vs. influent TDS

(Leitner and Associates, 1997) 29

Figure 3-6: Reverse osmosis water recovery vs. influent TDS

(Combined data) 30

Figure 3-7: Reverse osmosis capital cost vs. feed flow 31

Figure 3-8: Reverse osmosis capital cost vs. feed flow 32

Figure 3-9: Reverse osmosis electrical use vs. feed TDS 34

Figure 3-10: Sulfuric acid required to consume alkalinity 35

Figure 3-11: Ion exchange schematic and service cycle 41

Figure 3-12: Influent cation concentration vs. run time 46

Figure 3-13: Influent cation concentration vs. resin volume 47

Figure 3-14: Water recovery vs. influent cation concentration 49

Figure 3-15: Ion exchange water recovery vs. influent cation concentration 50

Figure 3-16: Ion exchange water recovery vs. cation removal 52

Figure 3-17: Ion exchange capital cost as a function of resin volume 53

Figure 3-18: Annual labor cost as a function of resin volume 56

Figure 3-19: Electrodialysis Process 57
x

Figure Page

Figure 3-20: Water recovery vs. TDS removed 62

Figure 3-21: EDR capital cost as a function of feed water flow rate and TDS 63

Figure 3-22: Electrical usage vs. TDS removed 65

Figure 3-23: Chemical and parts replacement vs. feed water flow rate 66

Figure 3-24: Schematic for HEEPM 67

Figure 3-25: HEEPM capital cost vs. feed water flow rate 69

Figure 3-26: HEEPM electrical usage vs. feed TDS 71

Figure 3-27: HEEPM acid required vs. feed TDS 73

Figure 4-1: Vertical tube falling film vapor compression evaporator 81

Figure 4-2: Evaporation Crystallization process for brine

A) Cbr-TDS < 200,000 mg/L B) Cbr-TDS  200,000 mg/L 83

Figure 4-3: Evaporator capital cost vs. brine flow rate 85

Figure 4-4: Energy required to concentrate brine to 20% solids 86

Figure 4-5: Crystallizer capital cost vs. concentrate flow rate 87

Figure 4-6: Energy required to crystallize/concentrate brine to 50% solids 89

Figure 4-7: Bipolar electrodialysis 90

Figure 4-8: Membrane distillation 91

Figure 5-1: Influent and Effluent water characteristics questions and toolbox flow 95

Figure 6-1: Reverse osmosis sensitivity results

A) Perturbation Plot B) Normal Plot 104

Figure 6-2: Ion exchange sensitivity results

A) Perturbation Plot B) Normal Plot 107
xi

Figure Page

Figure 6-3: Electrodialysis reversal sensitivity results

A) Perturbation Plot B) Normal Plot 109

Figure 6-4: HEEPM sensitivity results

A) Perturbation Plot B) Normal Plot 111

Figure 7-1: Scenario 1 treatment cost vs. sodium removal 115

Figure 7-2: Scenario 1 itemized treatment costs 116

Figure 7-3: Scenario 2 treatment cost vs. sodium removal 117

Figure 7-4: Scenario 2 itemized treatment costs 118

Figure 7-5: Scenario 3 treatment cost vs. sodium removal 119

Figure 7-6: Scenario 3 itemized treatment costs 120

Figure 7-7: Scenario 4 treatment cost vs. sodium removal 122

Figure 7-8: Scenario 4 itemized treatment costs 123

Figure 7-9: Scenario 5 treatment cost vs. sodium removal 124

Figure 7-10: Scenario 5 itemized treatment costs 125

Figure 7-11: Scenario 6 treatment cost vs. sodium removal 127

Figure 7-12: Scenario 7 treatment cost vs. sodium removal 128

Figure 7-13: Scenario 6 itemized treatment costs 128

Figure 7-14: Scenario 7 itemized treatment costs 129

Figure ARO-1: Reverse osmosis water recovery vs. influent TDS 147

Figure ARO-2: Reverse osmosis electrical usage vs. feed TDS 148

Figure AED-1: EDR capital cost vs. feed flow rate TDS removal 1300 mg/L 152

Figure AED-2: EDR capital cost vs. feed flow rate 1300 < TDS removal 3000 mg/L 152
xii

Figure Page

Figure AED-3: EDR capital cost vs. feed flow rate TDS removal = 10,000 mg/L 153

Figure AH-1: HEEPM capital cost vs. system capacity and feed TDS (EET, 2006) 156

Figure AH-2: HEEPM capital cost vs. feed water flow rate (flow 238 bbl/day) 156

Figure AH-3: HEEPM capital cost vs. feed water flow rate (238 < flow 2,381) 157

Figure AH-4: HEEPM capital cost vs. feed water flow rate (2,381 < flow 23,810) 157

Figure AH-5: HEEPM capital cost vs. feed water flow rate (23,810 < flow) 158

Figure AH-6: HEEPM operating and maintenance costs (EET, 2006) 158
xiii

LIST OF TABLES

Table Page

Table 2-1: PRB CBM water constituents and concentrations by watershed 6

Table 2-2: Constituent limits for discharging PRB CBM waters 8

Table 2-3: Constituent limits for water uses 11

Table 3-1: Definitions of variables illustrated in Figure 3-1 13

Table 3-2: Demineralization technologies and pretreatments 14

Table 3-3: Reverse osmosis chemical rejection efficiencies 27

Table 3-4: Reverse osmosis system design programs 28

Table 3-5: Ion selectivity coefficients for Dow G-26 Resin 38

Table 3-6: Variables outlined in Figure 3-11 41

Table 3-7: HEEPM chemical rejection efficiencies 68

Table 5-1: Toolbox user interface 93

Table 5-2: Toolbox default economic parameters 97

Table 5-3: Toolbox default brine management economic parameters 99

Table 5-4: Step by step instructions for using the toolbox 100

Table 6-1: Factors used in sensitivity analysis with high and low values assigned 102

Table 6-2: Reverse osmosis factorial design and ANOVA results 105

Table 6-3: Ion exchange factorial design and ANOVA results 106

Table 6-4: Electrodialysis reversal factorial design and ANOVA results 108

Table 6-5: HEEPM factorial design and ANOVA results 110

Table 7-1: Scenarios used to evaluate technology costs 113

Table 7-2: Water parameters used to evaluate technology costs 113
xiv

Table Page

Table 7-3: Scenario 1 economic parameter percent contribution 116

Table 7-4: Scenario 2 economic parameter percent contribution 118

Table 7-5: Scenario 3 economic parameter percent contribution 120

Table 7-6: Scenario 4 economic parameter percent contribution 123

Table 7-7: Scenario 5 economic parameter percent contribution 126

Table 7-8 Scenario 6 economic parameter percent contribution 129

Table 7-9 Scenario 7 economic parameter percent contribution 130

Table 7-10: Total treatment cost for all scenarios at 243mg/L sodium removal 130

Table AP-1: Greensand capital cost estimate 145

Table ARO-1: Capital costs for RO with source 148

Table ARO-2: Variable definitions and values assumed for Equation ARO.4 150

Table ARO-3: Values used and calculated to develop Figure 3-9 and Equation 3.27 151

Table AIX-1: Ion exchange capital cost data 151

Table AED-1: EDR capital cost data Leitner and Associates (1997) 153

Table AED-2: EDR capital cost data Lawrence et al. (1998) 153

Table AED-3: EDR capital cost data Urano (1977) 154

Table AED-4: EDR capital cost data Larson and Leitner (1979) 154

Table AH-1: HEEPM capital cost estimates 155

Table AH-2: HEEPM operating and maintenance cost estimate 155

Table D-1: Reverse osmosis factorial design with responses 163

Table D-2: Ion exchange factorial design with responses 164

Table D-3: Electrodialysis reversal factorial design with responses 165
xv

Table Page

Table D-4: HEEPM factorial design with responses 166

Table D-5: Sensitivity analysis influent water chemistry 167

Table D-6: Electrodialysis reversal factorial design with responses (labor included) 168

Table D-7: EDR factorial design and ANOVA results with labor included 169
1

1. INTRODUCTION

1.1 Background

Coal bed methane (CBM) production in Wyoming’s Powder River Basin (PRB) has seen

much growth in the past twenty years with less than 100 wells in 1987 to over 13,600 wells in

2004. At this rate of growth it is estimated that there will be 20,900 wells by 2010 (Ruckelshaus,

2005). Finley and Goolsby (2000) estimated that there is 25.2 trillion cubic feet (TCF) of

recoverable CBM in the PRB. Of the total recoverable CBM only 1.35 TCF has been withdrawn

through May 2004 (Ruckelshaus, 2005). However, the concern is not the quantity of CBM being

withdrawn, but rather the water that is associated with it.

Waters associated with CBM are generally termed CBM co-produced waters. These

waters must be withdrawn from the coal seam to desorb methane from the surface of the coal,

thus allowing the methane to be removed (Western Resources, 2005). PRB CBM co-produced

water production through May 2004 totaled 2.9 billion barrels with an estimated total water

production of 1852 billion barrels (Ruckelshaus, 2005). The water which is a by-product of CBM

extraction then must be managed in a manner that will not adversely impact the environment.

Currently CBM co-produced waters are managed in five ways. These include: surface

discharge to rivers or infiltration ponds, irrigation, livestock watering, atomization and road

application for dust control (WYDEQ, 2007). None of these options return the water to the

source aquifer. The primary concern is what is in the water.

PRB CBM waters contain excess sodium, and may be classified as being saline-sodic

(water with a sodium adsorption ratio (SAR) greater than 12 and electrical conductivity (EC)
2

greater than 4dS/m) (Keith et al., 2003). Waters with a SAR above 13 have the  to cause

soil to be hard and cloddy when dry, to crust badly, and to take water very slowly (Soltanpour

and Follett, 1999) which makes the soil unusable or less suitable for growing crops. These waters

also have the ability to alter ecosystems (Sterns et al., 2005) when managed with current

methods.

These concerns have caused regulations to be imposed by the Wyoming Department of

Environmental Quality and United States Environmental Protection Agency regarding the

discharge of CBM co-produced waters in the PRB. Regulations have made it necessary for CBM

waters to be treated before discharge if the water does not meet the discharge requirements

(WYDEQ, 2007).

Processes that have the ability to treat these waters to meet discharge requirements

related to elevated sodium concentration include: pressure driven membrane processes, electrical

processes, ion exchange processes, and thermal processes (Arthur et al., 2005). Currently,

reverse osmosis, a pressure driven membrane process, and ion exchange processes are used in

the PRB to treat CBM waters (WYDEQ, 2007). These processes do treat these waters to meet

discharge requirements (WYDEQ, 2007), but other processes and technologies may treat these

waters more economically.

Yet another issue is how to dispose of waste products that are generated as a result of

treating CBM produced waters. Currently, waste is either disposed of via deep well injection or

through the use of evaporation ponds. Deep well injection requires the use of injection wells

which are costly to install and have a set capacity (Mickley, 2001). In addition, the location of

these wells is determined based on site geology to ensure drinking water sources do not become

contaminated. The other technique currently used is evaporation ponds. A major issue
3

encountered with evaporation ponds is that the salts accumulate in the pond and the pond will

eventually reach a capacity at which time the dried solids must be disposed of in some manner.

These issues make it necessary to investigate other brine disposal techniques.

Currently no tool exists to compare the many technologies available to treat produced

waters and wastes generated on an equal platform. Such a tool would make it possible to decide

how best to treat a water based on the influent water characteristics, effluent water requirements,

and the abilities of the technologies capable of removing sodium from water.

1.2 Objectives

The objectives of this work were to use publicly available data to: (1) determine what

technologies can treat CBM co-produced waters to meet discharge and use requirements based

on the characteristics of the water and; (2) to develop a toolbox which will aid producers and

regulators when deciding how best to treat CBM produced waters based on the abilities and costs

associated with each technology.

Specific tasks were to:

(a) Characterize CBM co-produced waters in the PRB based on constituent concentrations,

including SAR, EC, and total dissolved solids (TDS).

(b) Identify possible uses for this resource and the required water characteristics for each use.

(c) Identify technologies that can be used to treat CBM water to meet the use requirements.

(d) Investigate each technology in detail and develop an understanding of the technologies

with regards to chemical rejection efficiency, pretreatment requirements, waste produced,

auxiliary chemical usage, and energy requirements.

(e) Identify and investigate treatment waste disposal technologies.
4

(f) Identify capital and operating costs associated with each technology.

(g) Construct and validate a “toolbox” model that incorporates all of the above criteria, thus

allowing one to find the best way to treat CBM water when compared to other

technologies on an equal platform.

(h) Perform a sensitivity analysis on the toolbox to identify key variables.
5

2. WATER CHARACTERIZATION

2.1 Produced Water Characteristics.

Much work has been done to determine the characteristics of PRB CBM waters (Rice et

al., 2000; McBeth et al., 2003; Patz et al., 2004; Patz et al., 2006; Jackson and Reddy, 2007). The

conclusion from research has shown that CBM water constituent concentrations vary from

watershed to watershed in the PRB (McBeth et al., 2003). There are five primary watersheds in

the PRB. The rivers that comprise these watersheds are the Cheyenne River, Belle Fourche

River, Little Powder River, Powder River and Tongue River.

CBM water data for the five watersheds are outlined in Table 2-1. Data have been

compiled from numerous studies and databases to provide a range of values for constituents in

CBM waters for each watershed. Water constituent data collection points were located on a

USGS map which has the various watersheds of the PRB delineated on it. High, low, and mean

water constituent concentrations where then determined for each watershed. Additional

information about water constituents in PRB CBM waters is available from the Wyoming State

Geologic Survey (http://ims.wrds.uwyo.edu/prb/prb.html). The Wyoming State Geologic Survey

has done much research related to CBM water constituents in the PRB. Data available at the

above link represents water constituents on a by-coal-formation basis.
Table 2-1: PRB CBM water constituents and concentrations by watershed. Data are adapted from Rice et al. (20

noted.

Constituent Units Cheyenne River Belle Fourche River Little Powder River Powder River 

Mean High Low Mean High Low Mean High Low Mean High Low 

Temperature oC 16.8 17.7 15.9 20.2 28.7 13.8 18.0 26.5 13.9 19.5 24.8 11.7

Specific Conductance μS/cm 705 770 640 1023 1640 470 1426 1660 1070 2097 3020 860 

Total dissolved solids mg/L 430 470 390 666 1120 270 966 1260 720 1411 2010 540 F - μg/L 1600 1600 1600 967 1700 500 760 1000 500 710 1400 420Cl– mg/L 8.2 10.0 6.3 14.7 64.0 5.2 9.9 12.0 8.9 11.7 18.0 6.4SO42– mg/L 8.9 17.0 0.7 2.5 12.0 0.0 1.2 3.0 0.1 1.5 8.1 0.1Br μg/L 65 90 40 144 85 20 72 80 60 76 11 50Alkalinity (as CaCO3) mg/L 465 490 440 742 1170 290 1112 1520 810 1585 2320 580 NH4+ mg/L 2.4 2.4 2.4 2.1 3.2 1.1 3.4 5.3 2.1 2.7 4.8 1.1Ca2+ mg/L 25.5 32.0 19.0 34.4 59.0 5.9 47.6 69.0 30.0 20.3 35.0 9.1 K mg/L 6.6 7.3 5.9 8.1 14.0 3.8 11.7 15.0 8.1 8.4 18.0 3.9Mg2+ mg/L 11 13 10 15 25 2 27 46 14 14 28 5 Na+ mg/L 140 150 130 219 390 110 308 360 240 563 800 220 Ba2+ μg/L 300 370 230 605 1300 140 988 1600 470 539 750 240 Fe μg/L 1500 2400 580 484 490 120 554 820 270 569 2800 20 Si mg/L 4.8 5.3 4.3 4.6 5.8 3.8 4.5 5.6 3.7 5.2 7.1 4.6Sr μg/L 455 560 350 763 1800 100 1054 1900 450 415 840 250aAl μg/L 6.7 5.4 11 aAs μg/L 2.3 1.5 4.7 aSe μg/L 0.71 0.75 0.83 aB μg/L 51 62 89 aCu μg/L 10 13 24 aMo μg/L 0.31 0.25 2.51 aMn μg/L 11 10.6 29 aCr μg/L 33.1 43.6 83.9 aZn μg/L 4.5 6.2 26.9 pH 7.1 7.1 7.1 7.2 7.5 6.9 7.2 7.6 6.8 7.5 7.7 7.4 SAR 5.9 6.0 5.7 8.0 12.0 6.1 9.1 11.0 7.7 23.6 32.0 13.0 aData are adapted from McBeth et al. (2003)bData are adapted from Ruckelshaus Institute (2005)cData are adapted from Jackson and Reddy (2007)
7

Based on reviewed data, the primary constituents present in PRB CBM waters are sodium

and alkalinity. Sodium and alkalinity concentrations can range from 110 to 800 mg/L and 290 to

2320 mg/L as CaCO3, respectively. The source of alkalinity is most likely bicarbonate because

bicarbonate is the primary source of alkalinity in the reported pH ranges 6.9 to 7.7 (Sawyer et al.,

1994).

2.2 Water Use Requirements

2.2.1 Surface Discharge

PRB CBM waters are most commonly discharged to a river or infiltration pond (Lamb,

2006). These waters can also be put to other beneficial uses if necessary. Water uses can include:

discharge, drinking water, livestock watering, irrigation, enhanced oil recovery, and cooling

water for electric generation. Each use has specific guidelines which may be followed, and/or

regulations which must be followed.

Discharge to surface water is regulated by the Wyoming Department of Environmental

Quality (WDEQ). If a CBM water does not meet discharge requirements it must be treated to

meet discharge requirements, or be managed another way (Lamb, 2006). A range of discharge

limits are outlined in Table 2-2. Data used to develop concentration ranges for constituent

discharge, were found by reviewing numerous Wyoming Pollutant Discharge Elimination

System (WYPDES) permits (WYPDES, 2006) (see Appendix C for a list of WYPDES permits

reviewed).

Discharge points from the WYPDES permits were located on a USGS map which has the

various watersheds of the PRB delineated on it. The data collected from the WYPDES permits
8

did not appear to have significant differences from watershed to watershed. However, new

watershed based permitting requirements are in the process of becoming law (WDEQ, 2007).

Table 2-2: Constituent Limits for Discharging PRB CBM waters (WYPDES, 2006).

Constituent Units High

1 Low2

Specific Conductance S/cm 7500 465

Total Dissolved Solids mg/L 5000 310

F

-

μg/L 4000 2000

Cl

- mg/L 230 46

SO4

2- mg/L 3000 412

NH4

+ mg/L 5 0.5

Na

+ mg/L 350 270

Ba

2+

μg/L 1800 1800

Fe μg/L 1000 125

Al μg/L 750 490

As μg/L 150 3.1

Cu μg/L 13.2 8.6

Mn μg/L 820 629

Zn μg/L 97 90

Radium 226 pCi/l 60 1

pH 9 6.5

SAR 24 10

1High represents the highest value permitted

2Low represents the lowest value permitted.

2.2.2 Other Uses

PRB CBM waters may be used for drinking water (municipal or bottled). The United

States Environmental Protection Agency (USEPA) has set standards for some constituent

concentrations if a water is used as a source of drinking water. The Maximum Contaminant

Level (MCL) is the highest allowable constituent concentration for water and is enforced by

federal law. A secondary MCL is a guideline set by the USEPA that is not enforced by law. The
9

World Health Organization (WHO) has also set guidelines for drinking water (WHO

recommended standards) which are not enforced by law. Both secondary MCLs and WHO

recommended standards are guidelines that prevent aesthetic or odor problems from being

present in drinking water (USEPA, 2006; WHO, 2006). Drinking water standards and

recommendations are outlined in Table 2-3.

PRB CBM waters have and are being used for livestock watering and irrigation.

Recommended constituent concentrations for both of these uses have been developed by Ayers

and Westcot (1994). A key issue with using PRB CBM waters for irrigation is Sodium

Adsorption Ratios (SAR) above the recommendations in Table 2-3 (Lamb, 2006). SAR values

above the recommendations may cause clay expansion of soils, making soils unusable to grow

crops (Ayers and Westcot, 1994). SAR is calculated with Equation 2.1 (Frank et al., 2006).

2

Ca Mg

Na

SAR

2+ 2+

+

+

= (2.1)

Where, Na

+

is the sodium concentration (mol/L), Ca

2+

is the calcium concentration (mol/L), and

Mg

2+

is the magnesium concentration (mol/L).

Enhanced oil recovery (EOR) and cooling water for energy generation are two other

possible uses for PRB CBM waters. Enhanced oil recovery is the process of pumping water into

an underground geologic feature that contains oil for the purpose of recovering oil. The water

requirements for this use have been developed by Royce et al. (1984). Royce et al. (1984) found

that the water characteristics for EOR are a range of TDS concentrations and total hardness

concentrations between 10,000 – 150,000 mg/L and 1 – 5,000 mg/L, respectively. Xie (2007)

who is researching the use of PRB CBM waters for EOR stated that the water characteristics for
10

EOR are flexible and depend primarily on the geologic formation from which oil will be

extracted. Xie (2007) recommended TDS concentrations be less than 4000 mg/L for EOR.

CBM waters may also be used as cooling water for energy generation. Very pure water

(TDS < 1) is required for this use to prevent corrosion and precipitate formation (Bueker,

1997).The required water chemistry for this use is outlined in Table 2-3. These data have been

adapted from Bueker (1997). PRB CBM waters are currently not being used as cooling water.
11

Table 2-3: Constituent limits for water uses

Drinking water Livestockc Irrigationc 

Constituent Units Higha Lowb H

Specific Conductance S/cm 1661* 801* 5000 3000 5

Total Dissolved Solids mg/L 500 2000 2000 

Dissolved Fluoride μg/L 4000 2000Chlorides mg/L 250b 1064 0Sulfates mg/L 250b 1920 0Br (regulated BrO3) μg/L 10 10Alkalinity (as CaCO3) mg/L 216* 104* 510 Nitrate as N mg/L 10 10 100 10Nitrite as N mg/L 1 1 10Ca2+ mg/L 300* 30* 800 K mg/L 2Mg2+ mg/L 300* 10* 400 122 Na (Dissolved Sodium) mg/L 200b 4370 920 Ba (Total Barium) μg/L 2000 2000 100Fe (Dissolved Iron) μg/L 300b Si mg/L Sr mg/LAl μg/L 200b 5000As μg/L 10 0 200Se μg/L 50 50 50B μg/L 5000 2000Cu μg/L 1300 1000 500 Mo μg/LMn μg/L 50b 50Cr μg/L 100 100 1000Zn μg/L 5000* 2400Total Radium 226 pCi/l 5 5pH 8.5* 6.5* 8.5 - 6 1Sodium Adsorption Ratio (SAR) 15a Constituent value is USEPA MCL, unless otherwise notedb Constituent value is either USEPA Secondary Standard or WHO recommended standard, unless otherwise notedcAyers and Westcot (1994)d Bueker (1997)

* Recommendation or solubility limit
12

3. TREATMENT TECHNOLOGIES

3.1 Overview

Analysis of CBM waters and the current use requirements show that the primary

constituent that must be removed is sodium. Sodium removal from water is difficult because

nearly every other constituent must be removed from the water before sodium is removed. Water

treatment technologies used to do this are generally termed demineralization technologies.

A number of demineralization technologies can treat CBM produced waters to remove

sodium. Technologies that can treat these waters include: electrodialysis (ED), electrodialysis

reversal (EDR), high efficiency electro-pressure membranes (HEEPMTM), electrodeionization,

capacitive deionization, advanced zeolites, constructed wetlands, rapid spray evaporation, and

several reverse osmosis (RO) and ion exchange (IX) configurations. Currently, only two of these

technologies, reverse osmosis and ion exchange, are being used to treat PRB CBM water. Four

of these demineralization technologies (IX, RO, EDR, and HEEPM) were investigated in detail.

3.1.1 Process

The general process used to treat CBM waters with selected demineralization

technologies is illustrated in Figure 3-1. CBM water is degassed to remove methane when it

enters the water treatment facility. The water is then stored in an equalization pond or tank

(Beagle, 2007; Drake, 2007; Young, 2006). Water is then pumped from the equalization pond

and split into a treatment stream and a bypass stream. This is necessary because most

demineralization technologies remove more constituents from the water than is necessary for the

water’s intended use.
13

Demineralization

Technology

Qi

Ci

Qe

Ce

Qby

Ci

Qt

Ct

Qf

Cf

Qbr

Cbr

Qp

Ci

Qpr

Cpr

Equalization

Storage

Equalization

Storage

Pretreatment

Iron Alk.

+

Methane

Removal

Source

Water

Figure 3-1: General treatment flow schematic for demineralization technologies. Iron and Alk.

represent iron and alkalinity removal pretreatments, respectively. + signifies chemical addition.

Other variables are defined in Table 3-1.

Table 3-1: Definition and units for variables illustrated in Figure 3-1. Units presented are

consistent with those used by industry. Unit conversions are in appendix E.

Variable Definition Units

Qi Water flow rate into facility bbl/day

Ci Constituent concentration into facility mg/L

Qby Bypass water flow rate bbl/day

Qp Pretreatment water flow rate bbl/day

Qf Feed water flow rate bbl/day

Cf Feed constituent concentration mg/L

Qt Treated water flow rate bbl/day

Ct Treated constituent concentration mg/L

Qpr Pretreatment residual flow rate bbl/day

Cpr Pretreatment residual concentration mg/L

Qbr Brine flow rate bbl/day

Cbr Brine concentration mg/L

Qe Effluent water flow rate bbl/day

Ce Effluent constituent concentration mg/L

The treatment stream is first pretreated to remove constituents that may be harmful to the

demineralization technology. Iron and/or alkalinity are two constituents that may be removed

during pretreatment. Table 3-2 outlines different pretreatments that may be required for selected

demineralization technologies.
14

Table 3-2: Demineralization technologies and pretreatments that may be required for selected

technologies.

Technology Methane Iron Alkalinity

Ion Exchange x

Reverse Osmosis x x x

Electrodialysis x x

HEEPM x x x

After pretreatment the water is passed through the demineralization technology where

constituents are removed based on their chemical rejection efficiency. The amount of a

constituent removed from the water by a treatment technology is referred to as the technology’s

chemical rejection efficiency. Chemical rejection efficiency () varies for different constituents

and is described by Equation 3.1 for a single constituent (Metcalf and Eddy, 2003).

f

f t

C

C - C

 = (3.1)

The treated water concentration (Ct,i) for a constituent is equal to the feed concentration (Cf,i) of

the constituent multiplied by one minus the chemical rejection efficiency for that constituent.

The treated water concentration is used to determine the final treated water fraction (), Equation

3.2. Water that is not treated constitutes the bypass stream.

e

t

i t

i e

Q

Q

C C

C C

 =

−

−

= (3.2)

The constituents that are removed from the water enter a brine stream, which contains

water as well as the constituents that were removed. The relative quantity of water that leaves the

demineralization technology is termed water recovery, and is different for different technologies.

The percent water recovery (r) for a demineralization technology is described by Equation 3.3

(Taylor and Jacobs, 1996).
15

100%

Q

Q

f

t ηr = × (3.3)

By performing a mass balance on the treatment stream the brine flow rate (Qbr),

constituent concentrations in the brine (Cbr), and feed flow rate (Qf) can be determined. The feed

flow rate, brine flow rate, and constituent concentrations in the brine are described by Equations

3.4, 3.5, and 3.6.

( ) ( )

1

r

r

r

r i

br



1 

1



1   Q

Q

−















 − ⋅

⋅ + 













 − ⋅ ⋅

=

β

(3.4)

[ ( )]

br

f f f f br

br,i Q

Q C C Q Q

C

⋅ − ⋅ −

= (3.5)

( ) i br

r

f Q Q





Q = − (3.6)

The treated water is blended with bypass water to meet use requirements after passing through

the demineralization technology.

3.1.2 Costs

Costs have been estimated for selected technologies and pretreatments. The source of

capital, operating and maintenance costs are described with their respective technologies. The

estimated costs are not exact, and should be considered accurate to within ± 30-50% of the actual

costs associated with each technology. This is the typical accuracy range of engineering cost

estimating (Westney, 1997).

The reported costs for the technologies are based on the feed flow rate (Qf), and are

reported in dollars per barrel of influent water. These costs can be adjusted to accommodate

blending with Equation 3.7.
16

i

f

i f Q

Q

bbl as Q

Cost

bbl as Q

Cost

⋅

















= (3.7)

3.2 Pretreatment

3.2.1 Methane removal

Methane removal from CBM water is necessary because the presence of methane in the

water poses a safety hazard due to the possibility of methane exploding or igniting while in the

water treatment system. Methane may be removed from water by aerating the water as it enters

equalization storage. Aeration may be achieved through the use of a waterfall, rapids, pumping

air into the water or a degasser. The technique used to remove methane from water is a matter of

preference. However, the most widely used method in the PRB is the degasser (Lamb, 2006). A

degasser is illustrated in Figure 3-2.

The degasser allows methane in the water to be captured and sold instead of being

emitted to the atmosphere. The other methods will work, but the extent of methane removal

should be assessed on a case by case basis. Also, for increased worker safety the methane levels

in the water should be regularly monitored to prevent ignition of methane. Methane may ignite at

concentrations as low as 5% by volume in air (Praxair, 2004). The cost of methane degassing

will vary depending on the method selected.
17

Untreated produced water

Water Distributor

MethaneMethane

Degassed Water

Methane to storage

Untreated produced water

Water Distributor

MethaneMethane

Degassed Water

Methane to storage

Figure 3-2: Schematic of a degasser used to remove methane from produced water.

3.2.2 Equalization Storage

Equalization storage is required for all technologies. Equalization storage allows for

relatively continuous treatment by providing a uniform water flow rate and relatively consistent

influent water concentrations to the demineralization technology. Blend water and water to be

treated is stored in the equalization pond.

Equalization storage volume is dependent on the water flow rate to the facility, and is

assumed to be equal to the daily water volume to the facility (Beagle, 2007). This allows for

continuous water treatment if CBM wells are temporarily shut down or other problems are

encountered.

Costs associated with equalization storage are excavation and pond liner cost (if a pond is

used), or the cost of storage tanks (if a pond is not used). Excavation costs are dependent on the

type of soil being excavated and the location of excavation. Pond liner costs are determined by
18

the thickness of the liner. Liner thicknesses typically range between 20 and 120 mil (Mickley,

2001). The unit cost for liner is approximately $0.01 per mil thickness per square foot of liner

(Mickley, 2001). The total cost for an equalization pond may range from $0.50/bbl to greater

than $6.38/bbl (Mickley, 2001). Water storage tank costs can range from $26.7/bbl

(Watertanks.com, 2007) to $54/bbl (Aqua Store, 2007) depending on the tank material and

dimensions.

3.2.3 Iron Removal

Iron will foul membranes and can not be removed from the membranes with acid

cleaning. This increases the membrane replacement frequency with a concurrent increase in

operating and maintenance costs (Bates, 1998). Therefore, to maintain membrane life, membrane

technologies such as reverse osmosis and electrodialysis require the removal of iron prior to

treatment if the influent iron concentration (Ci-Fe) exceeds 300μg/L (Coulter and Thomas, 1998).

Iron may be removed by: aeration with settling or filtration, chemical oxidation with

filtration, oxidizing filters (manganese greensand), or ion exchange (Land, 1999). The

recommended technique for iron removal is aeration with settling followed by manganese

greensand filtration. This process is outlined in the following paragraph.

The water is first aerated by flowing over rocks or a water fall as it enters equalization

storage. Aeration will increase the oxygen content of the water and allow soluble, ferrous, iron to

be oxidized to non-soluble, ferric, iron. A small amount of ferric iron will then form ferric

hydroxide precipitate and settle out of solution in equalization storage (this process is slow and

may not remove sufficient iron, but it will reduce the chemical demand for the next step). The

water will then flow through a manganese greensand filter which will remove more iron.
19

Manganese greensand is a zeolite coated with manganese dioxide, MnO2. This medium

removes 99% of iron from solution by oxidizing ferrous iron to ferric iron with the oxygen

molecules that are attached to the manganese, Equation 3.8 and 3.9 (New Mexico State

University, 1999). Once the iron is oxidized ferric hydroxide, which is non-soluble, forms and is

filtered out of solution in the lower parts of the filter. The filter will also remove suspended

solids from water, further decreasing the possibility of membrane fouling.

Fe

2+

+ Z[MnOx]

n  Z[Mn2Ox]n-1

+ Fe

3+

(3.8)

Fe

3+

+ 3H2O  Fe(OH)3 + 3H

+

(3.9)

Where, Fe

2+

is ferrous iron, Z[MnOx] is the greensand zeolite with a manganese oxide species

attached to it before iron oxidation, Z[Mn2Ox] is the greensand zeolite after iron oxidation, Fe

3+

is ferric iron, Fe(OH)3 (s) is ferric hydroxide precipitate that is filtered out of solution, n is the

charge of the zeolite species before iron oxidation, and (n-1) is the charge of the zeolite species

after iron oxidation.

The oxidizing capacity of manganese greensand is approximately 454 to 778 grams iron

and/or manganese per square foot of filter bed area (Inversand, 2006). Once the oxidizing

capacity of the medium is reached, the medium is regenerated with potassium permanganate

(KMnO4). The theoretical amount of potassium permanganate required to regenerate the medium

is 1 and 2 grams per gram of iron and manganese removed, respectively (Inversand, 2006). The

potassium permanganate demand is calculated with Equation 3.10.

i-Fe i-Mn KMnO4 Demand = C + 2 ⋅C (3.10)

Where, KMnO4 Demand is the potassium permanganate demand (mg/L), Ci-Fe is the influent iron

concentration of the water (mg/L), and Ci-Mn is the influent manganese concentration (mg/L).
20

The actual capacity of the medium must be determined on a case by case basis because

there are other constituents present in the water that are being oxidized in addition to iron and

manganese. Therefore, assumptions have been made to determine the filter area and potassium

permanganate demand. Assumptions include: a filter capacity of 648 grams iron and manganese

per square foot filter bed area, pretreatment water flow rate of 4gpm/sq. ft, filter depth of 3ft, and

99% iron, manganese and arsenic removal efficiency (Inversand, 2006).

3.2.3.1 Iron Pretreatment Capital Costs

Capital costs and operating and maintenance costs associated with the use of greensand to

remove iron from produced waters have been determined by using cost data developed by

Hoffman et al. (2006). Hoffman et al. (2006) developed a design manual for greensand filters

which contains an estimate with itemized costs for a system treating 500,000gpd. How these

costs have been adjusted for different flow rates and waters are discussed in the following

paragraphs.

Hoffman et al. (2006) determined capital costs for iron removal with greensand.

Components of the capital costs were itemized and included: greensand filtration media, filtering

vessels, backwash holding tanks, structure costs and miscellaneous equipment costs. How

individual capital costs were determined is outlined in Appendix A-P. A relationship between

pretreated water flow rate and capital cost, Figure 3-3, was developed from this analysis.

Equation 3.11 calculates the capital cost (CCpretreat) for iron removal from water by greensand

filtration.

0.5971

CCpretreat 1279 Qp = ⋅ (3.11)
21

y = 1279x

0.5971

R

2

= 0.9984

$0

$500,000

$1,000,000

$1,500,000

$2,000,000

$2,500,000

0 50,000 100,000 150,000 200,000 250,000

Pretreatment Water Flow Rate, bbl/day

C

a

pital C

ost, $

Figure 3-3: Greensand iron removal capital costs as a function of pretreatment water flow rate.

Adapted from Hoffman et al. (2006).

3.2.3.2 Iron Pretreatment Operating and Maintenance Costs

Operating and maintenance costs for iron pretreatment, Equation 3.12, include:

electricity, potassium permanganate, greensand replacement, and labor.

OCpre = OCelec-pre + OCchem−pre + OCmedia -pre + OClabor-pre(3.12)

Where, OCpre is the operating cost for iron removal pretreatment with manganese greensand

($/bbl), OCelec-pre is the electrical cost ($/bbl), OCchem.-pre is the chemical cost due to potassium

permanganate use ($/bbl), OCmedia-pre is the cost for greensand media replacement ($/bbl), and

OClabor-pre is the labor cost ($/bbl).

Operating costs due to electrical use have been developed from Hoffman et al. (2006).

The electrical use is assumed to be 0.0008 kW-hr/bbl of water pretreated (Equation 3.13).

elec pre C elec OC = 0.0008 ⋅ ost−

(3.13)
22

The cost associated with potassium permanganate (KMnO4) demand is dependent on the

iron and manganese concentration of the produced water, and the cost of potassium

permanganate. The demand is found with Equation 3.10. The operating and maintenance cost

associated with potassium permanganate is then the demand multiplied by the unit cost of

potassium permanganate, $3.80/lb (Chemistry Store, 2007), and conversion factors (Equation

3.14).













⋅

















⋅





















= ⋅

−

bbl

159 L

1,000 mg

1gram

lb

454 grams

lb

$3.80

OCchem pre KMnO4 Demand (3.14)

Greensand is assumed to need replaced every ten years. The cost associated with

greensand replacement is considered constant and equal to $0.0002/bbl (Hoffman, et al., 2006).

Labor is the final cost associated with iron removal. Hoffman et al. (2006) found that the

cost due to labor is $0.003/bbl. This is used to determine the operating and maintenance costs

associated with greensand pretreatment.

3.2.4 Acid Addition

Acid addition as a pretreatment may be necessary if a membrane technology (reverse

osmosis and/or electrodialysis) is being used. Waters containing large amounts of alkalinity have

the potential to foul membranes. Fouling occurs when calcium carbonate or other chemicals

precipitate and impede the flow of water through the membranes. Chemical precipitation occurs

due to a concentration gradient that forms between the membrane and the solution. The highest

concentrations form closest to the membrane and precipitates form if the concentration exceeds

the chemical’s water solubility. The solid precipitate then clogs the pore spaces of the membrane

which reduces the water flux through the membrane (Gwon et al., 2003). This phenomenon is

known as concentration polarization, and will decrease the operational life of the membranes.
23

Acid is typically added to the water to reduce the risk of concentration polarization occurring to

the degree that will cause undesirable affects (Metcalf and Eddy, 2003). Acids consume

alkalinity, therefore decreasing the possibility of precipitate formation.

The dominant contributor to alkalinity in PRB CBM waters is bicarbonate (HCO3

-

)

because the pH range for these waters is 6.8 to 7.7. The chemistry for bicarbonate consumption

by sulfuric acid (H2SO4) is illustrated in Equations 3.15 and 3.16.

2HCO3

-

+ H2SO4  2H2CO3 + SO4

2-

(3.15)

H2CO3 + H2O  CO2 + 2H2O (3.16)

The stoichiometry of Equation 3.15 shows that one mole of sulfuric acid consumes two moles of

bicarbonate. This allows for the approximate acid dose to be calculated for a specific water in

moles/L. Once the moles of acid required is known the mass of acid required can be calculated.

This procedure is outlined in Equation 3.17 for water with a bicarbonate concentration expressed

in units of mg/L and Equation 3.18 for water with a bicarbonate concentration expressed in units

of mg/L as calcium carbonate (CaCO3).

[ ]

































⋅

















= ⋅

2 4

2 4

3

2 4

2 4 3

mmol H SO

98 mg H SO

2 mol HCO

1mol H SO

61mg

1mmol

H SO Dose HCO (3.17)

g

mg

1000

L H SO

1,836 g H SO

36,700 meq

L H SO

50mg

1meq

H SO Dose Alk

2 4

2 4 2 4

2 4

⋅

































⋅

















= ⋅ (3.18)

Where, H2SO4 Dose is the sulfuric acid dose (mg/L), Alk is the bicarbonate concentration (mg/L

as CaCO3), 36700 and 1836 are the equivalent concentration (meq/L) and density (g/L) for 98%

pure sulfuric acid, respectively (DOW, 2005). Once the acid dose is known the cost of acid

pretreatment can be found by multiplying the acid dose by the unit price of acid.
24

Hydrochloric acid (HCl) or other acids may also be used. The hydrochloric acid dose has

been found as a function of the sulfuric acid dose by assuming sulfuric acid is 98% pure (EET,

2006) and hydrochloric acid is 36% pure (Beagle, 2007) (Equation 3.19).

HCl Dose 1.96 H SO Dose = × 2 4(3.19)

Hydrochloric acid may be used if calcium sulfate precipitation is a concern, although this

is not likely in PRB waters because of low calcium concentrations (much less than 100 mg/L).

Calcium sulfate is 136 times more soluble in water than calcium carbonate, 2.04 g/L for calcium

sulfate and 0.015 g/L for calcium carbonate at 20oC (Lenntech Water Treatment and Air

Purification, 2006). The chemistry of the water should be investigated before deciding which

acid to use.

The use of acid is not limited to pretreatment. Some technologies require the use of acid

to regenerate resins and clean technology specific components. Other uses of acid will be

discussed in the following sections.

3.3 Reverse Osmosis

3.3.1 Overview

Reverse osmosis (RO) is a demineralization technology that removes constituents from

water by forcing water through a semi-permeable membrane. This is done by exerting a pressure

on the water that is greater than the osmotic pressure of the water.

The osmotic pressure of a solution is related to the TDS concentration of the solution.

When two solutions having different TDS concentrations are separated by a semi-permeable

membrane, a difference in chemical potential will exist across the membrane. Water will diffuse

from the side of the membrane with lower TDS concentration to the side with higher TDS
25

concentration, attempting to equalize the chemical potential. This process is termed osmosis and

the flow from one side of the membrane to the other will continue until the chemical potential is

equal on each side of the membrane. The end result is a pressure differential between the two

sides of the membrane. This pressure differential is the osmotic pressure of the high TDS

solution and can be calculated with equation 3.20. The process of osmosis is illustrated in Figure

3-4a.

 = Cf,TDS⋅R ⋅T (3.20)

Where,  is the osmotic pressure of the solution (atm), Cf,TDS TDS concentration of the solution

(mol/L), R is the gas constant (0.08205 L atm mol

-1 K-1

) and T is the solution temperature (K).

Reverse osmosis is the opposite of osmosis. Reverse osmosis is simply applying a

pressure to the high TDS concentration side of the membrane that is equal to or greater than the

osmotic pressure. This causes water to flow across the membrane to the lower TDS side which

produces a lower TDS permeate, and a concentrated brine stream. This process is illustrated in

Figure 3-4b.

Figure 3-4: Reverse osmosis treatment process.  is the osmotic pressure of the high TDS

solution and 	P is an external pressure applied to the high TDS solution. a) For Osmosis 	P is

equal to atmospheric pressure. b) For Reverse osmosis 	P is a pressure greater than .

Low TDS High TDS

M

e

m

bra

n

e

Low TDS High TDS

M

e

m

bra

n

e

P

a) Osmosis: P < 

Brine

b) Reverse Osmosis: P > 

P


26

After pretreatment, the water is pressurized with high pressure pumps to a pressure

greater than the osmotic pressure of the water, 15-25 bar for brackish water and 54-80 bar for

seawater (Miller, 2003). This pressure is termed the feed pressure. The water is then forced

through a semi-permeable membrane that rejects a percentage of the salts in the feed solution.

Rejected salts are then removed from the system in a brine, and the cleaned water is discharged

or stored for other uses.

Wijmans and Baker (1995) have shown that RO is a diffusion limited process in which

the movement of solutes and water across the membrane may be described by the solutiondiffusion model, Equations 3.21 and 3.22, respectively.

J = A(∆P - ∆π ) i(3.21)

( ) j jo Cj

J = B C - (3.22)

Where, Ji and Jj are the water and solute flux, respectively (kg/m

2

s), A is the water permeability

constant which is membrane specific (s/m), P is the pressure applied to the feed side of the

membrane (kPa),  is the osmotic pressure difference across the membrane (kPa), B is the

solute permeability constant which is membrane and solute specific (m/s), Cjo is the solute

concentration on the feed side of the membrane and Cjis the solute concentration on the effluent

side of the membrane (kg/m

3

).

Upon investigation of the solution-diffusion model it becomes evident that the movement

of water through the membrane is dependent on the osmotic pressure of the feed solution, the

pressure applied and the rate at which water can diffuse through the membrane. However, the

movement of solutes such as salts through the membrane is dependent on the solute

concentration at the membrane surface on each side of the membrane and the membrane specific
27

solute diffusivity. Therefore, the chemical rejection efficiency, or percent of a specific

constituent removed from the water, is not 100%.

The chemical rejection efficiency for reverse osmosis depends primarily on the

membrane being used (Yoon and Lueptow, 2005; Dow, 2005). A range of chemical rejection

efficiencies are shown in Table 3-3. These chemical rejection efficiencies were found by using

ROSA, a reverse osmosis design program developed by Dow Chemical Company. This program

gives the user the option to select from several different Dow FILMTEC

TM membranes. Table 3-3

is the result of running ROSA ten different times using a different membrane each time. The

source water chemistry input was Powder River High (see Table 2-1) and the system was

designed with 85% water recovery. Average values from Table 3-3 are the assumed chemical

rejection efficiencies for RO.

Table 3-3: Reverse Osmosis chemical rejection efficiencies for various constituents. Data was

developed using ROSA software and is the average, high, and low values from ten different

FILMTEC

TM membranes operating at 85% recovery with a 4406mg/L feed TDS.

Constituent Average High Low

NH4

+

92% 98% 89%

K

+

96% 99% 93%

Na

+

96% 99% 93%

Mg

2+

97% 99% 95%

Ca

2+

97% 99% 96%

Sr

2+

97% 99% 96%

Ba

2+

97% 99% 96%

HCO3

-

95% 99% 93%

NO3

-

48% 75% 32%

Cl

-

96% 99% 76%

F

-

95% 99% 92%

SO4

-2

99% 100% 98%

SiO2 97% 99% 95%

Boron 28% 67% 11%

TDS 96% 99% 94%

Other 90%
28

3.3.2 Water Recovery

Water recovery for RO is a function of several variables. The type of membrane used,

feed pressure, number of membranes used in series (stages), and feed water characteristics all

have an impact on the water recovery. All of these variables are interrelated and the design of a

system is generally based on the desired product water recovery. This has caused many RO

membrane manufacturers to develop computer programs for the design of RO systems. A few of

these programs are listed in Table 3-4. The complexity involved in the design of an RO system is

beyond the scope of this work so another method for determining water recovery was utilized.

Table 3-4: Reverse Osmosis system design programs by manufacturer

Manufacturer Program

Hydranautics IMSDesign

GE/Osmonics WINFLOWSTM

Dow ROSA

IDS CAROL

Three sources of data were used to find a function relating TDS to water recovery. EET

Corporation provided a plot of water recovery versus influent TDS concentration (see Appendix

A-RO Figure ARO-1). This plot assumes a consistent feed pressure of 800psi and the same

effluent TDS concentration of 500mg/L. The validity of this plot is somewhat uncertain because

the data are from a plot comparing the recovery of HEEPM (a technology produced by EET

Corporation, see section 3.6) to an RO system.

Another source of data was Leitner and Associates (1997). In 1997 Leitner and

Associates, Inc. conducted a survey of several brackish water desalination plants in the United
29

States. Data collected included: capital cost, labor requirements, process chemical costs, influent

and effluent TDS concentrations, and water recovery. Figure 3-5 was developed from Leitner

and Associates (1997) to validate Figure ARO-1. The Leitner and Associates (1997) data do not

show a strong correlation between feed TDS and water recovery for the range of feed TDS

concentrations between 50 and 7000 mg/L. Because only one data point is present beyond 7000

mg/L the averages and standard deviations were found for feed TDS concentrations (1902 ±

1614 mg/L) and water recovery (79 ± 8%) for feed TDS concentrations between 50 and 7000

mg/L. The average and standard deviation are plotted with other data sources on Figure 3-6.

y = -0.0015x + 81.969

R

2

= 0.4678

0

10

20

30

40

50

60

70

80

90

100

0 5000 10000 15000 20000 25000

Feed TDS, mg/L

W

ater Reco

very, %

Figure 3-5: Reverse osmosis water recovery with respect to influent TDS for a variety of

effluent TDS concentrations. Adapted from Leitner and Associates (1997).

The final source of RO water recovery data is from the Petro-Canada Wild Turkey two

stage RO plant in the PRB. Wild Turkey is currently operated by Siemens Corp. During a visit to

the PRB, Jeff Young of Siemens informed us that the plant is currently treating water with an
30

influent TDS of 2400 mg/L at 95% water recovery (Young, 2006). These data in addition to the

other data sources are illustrated in Figure 3-6.

y = -0.0015x + 100

R

2

= 1

y = -0.0015x + 95.149

R

2

= 0.8354

0

10

20

30

40

50

60

70

80

90

100

0 5000 10000 15000 20000 25000 30000 35000 40000

Feed TDS, mg/L

W

ater Reco

very, %

EET, 2006

Young, 2006

Leitner and Assoc., 1997

Linear (EET, 2006)

Linear (Combined)

Figure 3-6: Reverse osmosis water recovery with respect to influent TDS. Combined refers to

the data set containing all data sources, which includes: EET (2006), Leitner and Associates

(1997), and Young (2006). EET (2006) data assume an influent feed pressure equal to 800psi for

a predominately NaCl feed solution while maintaining a constant effluent TDS concentration

equal to 500mg/L. The Leitner and Associates (1997) point with error bars corresponds to the

average and standard deviation of TDS values between 50 and 7000 mg/L.

EET (2006), Leitner and Associates (1997) and Young (2006) data, are illustrated in

Figure 3-6. The slopes of the lines in Figure 3-6 are the same, but the y-intercepts are different.

The y-intercept for the “combined” data is approximately 95% while that for the EET Corp. data

is 100%. The EET Corp. data are more realistic because: (a) water with a feed TDS equal to zero

would have a water recovery of 100% because no brine is being produced, and (b) Young (2006)

data agree well with EET (2006). Therefore, the EET Corp. data are used to calculate the water

recovery. The water recovery is calculated with Equation 3.23.

r 0 Cf -TDS  = 100 - .0015⋅ (3.23)
31

3.3.3 Capital Costs

RO capital costs have been determined with the use of data collected by International

Consultants et al. (2000) and Beardsley et al. (1994). International Consultants et al. (2000)

developed RO capital costs with the WATERCO$T Model and Water Model; while Beardsley et

al. (1994) attained estimates from US Filter and Glegg Water Conditioning. All capital cost data

collected from these sources have been adjusted to January 2007 U.S. dollars by using the

Engineering News-Record construction cost indices. These data are located in Appendix A-RO

and are illustrated in Figures 3-7 and 3-8.

y = 21.844x + 2E+06

R2 = 0.999

$0

$30,000,000

$60,000,000

$90,000,000

$120,000,000

0 1,000,000 2,000,000 3,000,000 4,000,000 5,000,000

C

a

pital C

ost

Feed Water Flow Rate, bbl/day

Figure 3-7: Data developed by International Consultants et al. (2000) with WATERCO$T

Model. This data is used to calculate RO capital costs for feed water flow rates greater than

54327 bbl/day.

RO capital costs are dependent on the feed water flow rate. Economies of scale become

apparent at a feed water flow rate equal to 54327 bbl/day based on the data presented. RO capital

costs are calculated with the functions on Figures 3-7 and 3-8, Equations 3.24 and 3.25.
32

CCRO = 56.799 ⋅Qf +101006 Qf 54327 bbl/day (3.24)

CCRO = 21.844 ⋅Qf + 2000000 Qf > 54327 bbl/day (3.25)

Where, CCRO is the capital cost of a reverse osmosis system ($) and Qf is the feed water flow rate

(bbl/day).

y = 56.799x + 101006

R2 = 0.9836

y = 21.844x + 2E+06

R2 = 0.999

$0

$1,000,000

$2,000,000

$3,000,000

$4,000,000

$5,000,000

$6,000,000

$7,000,000

$8,000,000

0 50,000 100,000 150,000 200,000 250,000

C

a

pital Cost

Feed Water Flow Rate, bbl/day

Beardsley et al., 1994

International Consultants et al., 2000 (WATERCO$T Model)

International Consultants et al., 2000 (Water Model)

Q < = 54327 bbl/day

Q > 54327 bbl/day

Figure 3-8: Reverse Osmosis Capital Costs. Adapted from International Consultants et al.

(2000) and Beardsley et al. (1994). Q > 54327 bbl/day is for feed water flows greater than 54327

bbl/day and Q < 54327 bbl/day is for feed water flows less than or equal to 54327 bbl/day.

3.3.4 Operating and Maintenance Costs

RO operating and maintenance costs include: electricity, labor, chemicals, and membrane

and miscellaneous parts replacement. The methods used to approximate these costs are outlined

below. The total operating and maintenance cost for RO is calculated using Equation 3.26.

OCRO = OCelec-RO + OClabor-RO + OCchem-RO + OCmem-RO + OCparts-RO (3.26)
33

Where, OCRO is the total operating and maintenance cost for RO ($/bbl), OCelec-RO is the

electrical cost ($/bbl), OClabor-RO is the cost due to labor ($/bbl), OCchem-RO is the chemical cost

($/bbl), OCmem-RO is the membrane replacement cost ($/bbl), and OCparts-RO is miscellaneous parts

replacement cost ($/bbl).

RO requires electricity for high pressure pumps that impart energy to the water. Taylor

and Jacobs (1996) have developed equations that relate the theoretical amount of energy required

to feed water TDS, permeate TDS, concentrate TDS, water recovery, membrane water

permeability, and other system parameters. This method was used to determine the electrical

requirement for a RO system. This method, assumptions, and equations are presented in

Appendix A-RO. The results are illustrated in Figure 3-9.

Data developed from the Taylor and Jacobs (1996) method was then verified by

comparing them to Leitner and Associates (1997) data (see Appendix A-RO Figure ARO-2).

This was done by finding the averages and standard deviations for the feed TDS (1627 ± 1471

mg/L) and electrical requirement (0.19 ± 0.07 kW-hr/bbl) from Leitner and Associates (1997).

Leitner and Associates (1997) averages and standard deviations are illustrated on Figure 3-9 with

error bars. Since the Leitner and Associates (1997) data are on the line developed using the

Taylor and Jacobs (1996) method, the equation developed from the Taylor and Jacobs (1996)

method is used to calculate the electrical requirement for RO. The cost of electricity for RO is

calculated with equation 3.27.

(( ) ) f,TDS elec

-5 OCelec-RO = 0.1753⋅ exp 3⋅10 ⋅C ⋅Cost (3.27)

Where, OCelec is the operating cost due to electricity ($/bbl), Cf,TDS is the feed TDS concentration

(mg/L), and Costelec is the unit cost of electricity ($/kW-hr).
34

y = 0.1753e

3E-05x

R

2

= 0.9971

0.0000

0.1000

0.2000

0.3000

0.4000

0.5000

0.6000

0 5000 10000 15000 20000 25000 30000 35000 40000

Feed TDS, mg/L

Taylor and Jacobs, 1996

Leitner and Assoc., 1997

Electrical Use, k

W-hr/b

bl

Figure 3-9: Electrical energy requirement with respect to feed TDS for reverse osmosis.

Developed using the Taylor and Jacobs (1996) method (see Appendix A-RO).

The labor requirement for a RO system has been developed from Leitner and Associates

(1997), and is assumed to be 1 full time employee per 11,682 bbl of feed water per day. The cost

for an employee (Costemployee) is assumed to be $50,000/yr. The cost of labor per barrel of feed

water (CLabor-RO) is calculated with Equation 3.28.

Labor-RO C employee

ost

365 day

1 yr

day

bbl

11,682

1employee

OC ⋅

















⋅

























= (3.28)

Equation 3.28 shows that OCLabor-RO is independent of the flow rate. Therefore, the cost of labor

is considered fixed and equal to $0.012/bbl.

Chemical usage is a function of the feed water characteristics. Chemicals that may be

required are acids and antiscalants. Acids may be necessary to consume alkalinity as described in
35

pretreatment. Antiscalants may be necessary to prevent further scale formation on the

membranes. Due to the proprietary nature of antiscalants only acid addition for the consumption

of alkalinity is considered.

The amount of acid required to consume all of the alkalinity in the feed water is

calculated using the methods developed in pretreatment. Figure 3-10 was developed by utilizing

Equations 3.17 and 3.18. The function developed and illustrated on Figure 3-10 is used to

determine the amount of sulfuric acid required to consume all of the alkalinity in the feed water.

Equation 3.29 is used to calculate the cost of acid per barrel of feed water.

y = 2E-07x

0.0000

0.0005

0.0010

0.0015

0.0020

0 2000 4000 6000 8000 10000 12000

Feed Alkalinity, mg/L as CaCO3

T

o

ns S

ulfuric Acid

per b

bl w

ater

Figure 3-10: Tons of 98% sulfuric required to consume all of the alkalinity in the feed water

( ) f,Alk acid

7 OCchem-RO = 2×10 ⋅C ⋅Cost−

(3.29)

Where, OCchem.-RO is the cost of acid per barrel of feed water ($/bbl), Cf,Alk is the feed water

alkalinity (mg/L as CaCO3) and Costacid is the unit cost of acid ($/Ton). This function can be
36

modified if 36% hydrochloric acid is used instead of 100% sulfuric acid by multiplying the result

by 1.96.

Membrane replacement and miscellaneous parts replacement were determined by Wade

(2001). The cost of membrane replacement and miscellaneous parts are $0.017/bbl and

$0.00525/bbl, respectively.

3.3.5 Energy Recovery and Savings

RO systems can be designed with energy recovery devices that can recover energy from

the water that would otherwise be wasted upon discharge. The primary method of energy

recovery is through the use of energy recovery turbines (ERT). An ERT is an electrical generator

that uses the energy of the high pressure water to turn an impeller. These turbines can recover

energy from the water after it has been treated. This is possible due to the water flowing at

velocities capable of turning a turbine after treatment.

Once the water is pressurized and pushed through the membrane the water maintains a

flow rate equal to the daily flow rate minus the brine flow rate. Darwish and Al-Najem (1999)

found that 444 kW of power can be recovered from a RO system that requires 1633 kW when

using ERTs to recover energy from only the brine stream. This correlates to 27% energy savings.

Due to the energy savings alone, the use of ERTs should be investigated when designing an RO

system. Costs associated with the use of ERTs have not been included due to these costs being

beyond the scope of this work.
37

3.4 Ion Exchange

3.4.1 Overview

Ion exchange is the process of exchanging ions that are on an insoluble exchange material

for ions that are dissolved in solution (Metcalf and Eddy, 2003). This process is an adsorption

phenomenon where the mechanism for adsorption is primarily electrostatic (Weber and DiGiano,

1996). Ions are held by electrostatic forces to charged functional groups on the surface of the ion

exchange material (AWWA, 1990). As water, which contains ions, contacts the resin, ions on the

resin are exchanged for ions in the water. This process has been used in water and wastewater

treatment for: demineralization, desalting, hardness removal, nutrient removal, and in the

recovery of valuable waste products (Reynolds, 1982; Weber and DiGiano, 1996).

The exchange material can be either naturally occurring, such as zeolites, soils, etc., or a

synthetic resin, and can remove either cations or anions from solution depending on the material

used (Reynolds, 1982). Cation exchange materials remove sodium from water and are therefore

the focus of this work. Equation 3.30 shows how this process works for a cation exchange

material that exchanges hydrogen ions on the resin for sodium ions in the water.

RH + Na

+  RNa + H+

(3.30)

Where, RH is a cation exchange material with a hydrogen ion attached to it and RNa is the

exchange material with a sodium ion attached to it after the water has contacted the exchange

material.

In the case depicted above hydrogen ions are released into the water. These hydrogen

ions will consume alkalinity and generate carbon dioxide as shown in Equations 3.15 and 3.16.

In this case it is necessary to remove carbon dioxide from the water if the resulting pH is too low

for discharge (Reynolds, 1982).
38

The previous example is true for water that contains only sodium and bicarbonate, but

there is more than just sodium in PRB CBM waters. The presence of other ions gives way to

competition among ions to adsorb to the surface of the exchange material.

The degree to which ion exchange occurs for a certain constituent is related to the ionic

strength of that constituent. This concept is best described by the ion selectivity coefficient.

Table 3-5 lists several ion selectivity coefficients for Dow G-26 strong acid cation exchange

resin. The selectivity coefficients are in order of the resins preference for a specific ion versus

hydrogen (which is the original ion on the resin). For example: the selectivity coefficient for

sodium vs. hydrogen is 1.56 while calcium has a selectivity coefficient of 4.06. As a result

calcium is selectively removed before sodium.

Table 3-5: Ion selectivity coefficients for Dow G-26 Resin (Dow, 2007)

Counter Ion Selectivity Coefficient

Li

+

0.79

H

+

1

Na

+

1.56

NH4

+

2.01

K

+

2.28

Rb

+

2.49

Cs

+

2.56

Mg

2+

2.59

Zn

2+

2.73

Co

2+

2.94

Cu

2+

3.03

Cd

2+

3.06

Ni

2+

3.09

Ca

2+

4.06

Sr

2+

5.13

Pb

2+

7.8

Ba

2+

9.06
39

3.4.2 Process

Ion exchange is generally performed in a vessel containing ion exchange material (here

on referred to as resin unless otherwise noted). Water contacts the ion exchange material at a

flow rate that allows for adequate contact time with the resin to achieve the desired effluent ion

concentration. The water is then blended with untreated water if excessive ions were removed,

and finally discharged or stored for other uses.

This process can continue until the ion exchange material reaches its capacity (exhaustion

or breakthrough). This is the point at which the ion exchange material no longer removes the

target ion, e.g. sodium, from the water. When breakthrough occurs the target ion becomes

present in the effluent water at concentrations greater than before breakthrough. At this time the

ion exchange material is either disposed of or regenerated.

Regeneration is the process of removing ions from the ion exchange material that were

removed from the water and restoring the original ions to the material (the opposite of Equation

3.30). Regeneration is performed by contacting the resin with a solution which contains the

original ion, regenerant (Reynolds, 1982). The regenerant typically contains a high (1-10% by

weight) concentration of the original ion to ensure that the resin is restored to its original state

(Wachinski, 2006).

For sodium removal, the regenerant solution must contain hydrogen ions because

hydrogen ions have a lower selectivity than sodium ions. Therefore, an acid is used to regenerate

the resin when sodium is the target ion (Dow, 2007). Regeneration is described by Equation 3.31

when using sulfuric acid to regenerate the material used in Equation 3.30.

2RNa + H2SO4  2RH + Na2SO4 (3.31)
40

Hydrochloric acid may also be used to regenerate this resin. This reaction is described by

Equation 3.32.

RNa + HCl  RH + NaCl (3.32)

On a molar basis more hydrochloric acid is required than sulfuric acid to regenerate the

same material when comparing Equations 3.31 to 3.32. However, ion exchange capacity is not

expressed in units of moles, but is rather expressed in units of equivalents per liter of resin.

On an equivalent basis more sulfuric acid is required than hydrochloric acid. This is because the

equivalent weight of hydrochloric acid is 36.5 g/eq while for sulfuric acid the equivalent weight

is 49 g/eq. This becomes important when assessing the waste stream generated from an ion

exchange process.

Overall the process of ion exchange consists of three steps which are carried out in the

following order:

1) Service: untreated water contacts the resin and the target ions are removed from the

water, and displace hydrogen ions on the resin.

2) Regeneration: an acid contacts the resin, and hydrogen ions displace the target ions from

the resin, thus restoring the resin to its original state.

3) Rinse: water is flushed through the system to remove residual regeneration chemicals.

Upon completion of the rinse cycle the process starts back at step 1. This process is outlined in

Figure 3-11 for a downward flow (co-current) service cycle and upward flow (counter-current)

regeneration cycle. Variables illustrated in Figure 3-11 are defined in Table 3-6.
41

Ion

Exchange

Resin

Ion

Exchange

Resin

Ion

Exchange

Resin

Ion

Exchange

Resin

QR

QR

Ion

Exchange

Resin

Ion

Exchange

Resin

Qr

Qr

Ion

Exchange

Resin

Ion

Exchange

Resin

Qf

Cf

Qf

Ct

Qf

Ct

1 2 3

Ct < Ctarget Ct Ctarget

CR-in Cr-out

CR-out

Regeneration Rinse

Cr-in

Qf

Cf

1 1

Figure 3-11: Schematic and service cycle for ion exchange process with co-current service cycle

and counter-current regeneration. Numbers and variables are described in text and Table 3-6,

respectively.

Table 3-6: Variables outlined in Figure 3-11

Variable Definition Units

Qf Feed water flow rate L/d

Cf Feed ion concentration meq/L

Ct Treated water ion concentration meq/L

QR Regenerant flow rate L/d

CR-in Influent regenerant concentration meq/L

CR-out Effluent regenerant concentration meq/L

Qr Rinse water flow rate L/d

Cr-in Influent rinse water concentration meq/L

Cr-out Effluent rinse water concentration meq/L

3.4.3 Design Requirements

A strong acid cation exchange resin is recommended for the removal of sodium from

PRB CBM waters (Dow, 2007). Subsequently, this is the type of resin currently being used in the

PRB by EMIT and Drake (Beagle, 2006; Drake, 2007). Dow Chemical Company recommends

their G-26 strong-acid cation exchange resin specifically for treatment of PRB CBM waters

(Dow, 2007). Therefore, Dow G-26 resin will be used in this analysis.
42

Design requirements for a cation exchange system depend on the resin capacity, influent

cation concentration, feed water flow rate, desired time between regenerations, and volume of

resin. The following equations are used to design an ion exchange system:

Resin

f, cat

water V

C

RC

V ≤ ⋅ (3.33)

f

water

Q

V


 = (3.34)

Qf = Avessel⋅ q (3.35)

RC

C

V 
 Q

f, cat

Resin f = ⋅ ⋅ (3.36)

Where, Vwater is the volume of water treated per service cycle (L), RC is the resin capacity (2000

meq/L of resin (Dow, 2007)), Cf, cat is the total influent cation concentration (meq/L), VResin is the

volume of resin that contacts the water (L), 
 is the run time or time between regenerations (d),

Qf is the feed water flow rate (L/d or m

3

/day), Avessel is the area of the vessel (m

2

), and q is the

service flow rate (m/h).

3.4.3.1 Regeneration

The theoretical amount of regenerant required to regenerate the resin is 1 equivalent

regenerant per 1 equivalent cation removed. Doug Beagle of EMIT Water Discharge

Technologies recommends 1.1 equivalents regenerant per equivalent of cation removed (Beagle,

2007). This will be used to determine the volume of regenerant required. The regenerant volume

required per regeneration is found using Equations 3.37 and 3.38.

regen

regen

EC

Regeneration

Cations Removed

1.1

Regeneration

V 















= ⋅ (3.37)
43













⋅

















= ⋅ ⋅ ⋅

bbl

42 gal

gal

3.78 L

C Q

Regeneration

Cations Removed

f,cat f

τ (3.38)

Where, Vregen is the volume of regenerant required per regeneration to restore the resin to its

original state (L), ECregen is the concentration of the regenerant solution (meq/L), and

CationsRemoved is the equivalents of cations removed from the water during a service cycle or

from the media per regeneration (meq), as shown in Equation 3.38.

After regeneration the resin must be rinsed to remove residual acid from the resin. Dow

Chemical Company recommends using 3 to 6 resin volumes of water per service cycle for

rinsing (Dow, 2007). The actual volume required will vary based on effluent requirements. Three

resin volumes of rinse water per service cycle is used in this analysis.

3.4.3.2 Waste stream

Two waste streams are generated from an ion exchange process. The first waste stream is

due to regeneration. As the regenerant contacts the exhausted resin hydrogen ions adsorb to the

resin and the ions that were on the resin enter the regenerant, as described above. The result is a

concentrated waste stream, brine. The volume of brine generated is a function of the type of acid

used to regenerate the brine, the concentration of the acid used, if water is reused, and most

importantly the influent cation concentration. The second waste stream is generated as a result of

rinsing the resin after regeneration. Both waste streams combined constitute the total waste

generated from the ion exchange process Equation 3.39.

Vbrine = Vregen + Vrinse(3.39)

Where Vbrine is the volume of brine generated per service cycle (L), Vregen is the volume of

regenerant used to regenerate the resin (L), and Vrinse is the volume of water used to rinse the

resin after regeneration (L).
44

3.4.4 Configurations

There are several different ion exchange configurations. The primary difference between

each configuration involves when and where the resin is regenerated, how the water contacts the

resin, and how the waste stream is managed.

3.4.4.1 Regeneration Configurations

There are two resin regeneration configurations of ion exchange, non-continuous bed and

continuous bed. Non-continuous bed ion exchange is a process where all the steps involved

occur in the same vessel at different times. This process is typically designed based on the

volume of resin required to achieve a certain runtime or time between regenerations (Wachinski,

2006). This process also requires an additional vessel that is capable of treating the water while

the other vessel is being regenerated and rinsed. These systems are typically less complex than a

continuous bed system.

Continuous bed ion exchange involves all steps being performed at the same time in

either separate vessels (Drake Process), or separate areas of the same vessel (Higgins loop). Flow

rate is the primary design variable because the resin volume is not dependent on the run time.

The result is a smaller treatment area and resin volume being required, but increased system

complexity (Wachinski, 2006).

There are currently two continuous bed configurations, the Higgins loop and Drake

process. The Drake Process involves the simultaneous movement of water and resin. The process

begins by dosing the water with a specific volume of resin. The resin dose and water-resin

contact time is determined by the required cation removal. Treated water and resin then move to

another vessel where they are separated and carbon dioxide is removed. The resin is then sent to

another vessel for regeneration and rinsing. After regeneration the resin is returned to the first
45

vessel to treat more water. This process is uninterrupted, making it continuous bed ion exchange

(Drake, 2007).

The Higgins loop continuous bed ion exchange system consists of pulsing or moving a

small portion of the resin bed out of the main column to another zone of the vessel for

regeneration and simultaneously moving a regenerated portion back to the column in which

water is being treated. The primary difference between the Higgins loop and Drake Process is

that the Drake Process doses the water with resin while the Higgins loop allows water to flow

through a resin bed.

3.4.4.2 Continuous bed vs. Non-continuous bed

To identify the primary differences and the reason why a non-continuous process may be

preferred over a continuous process the following analysis has been performed.

The run time for a non-continuous process can be described by Equation 3.34, and is a

function of the influent ion concentration or desired ion removal, the resin volume, feed water

flow rate, and resin capacity. These relationships are illustrated in Figure 3-12, when maintaining

a constant resin volume, and resin capacity. As the cation concentration increases the run time

decreases. This may result in a very short run time. This is not a problem for a continuous ion

exchange system such as the Higgins loop or Drake process in which the pulsing frequency

(Higgins Loop) or resin dose (Drake Process) would simply be adjusted. In a non-continuous bed

a very short run time may be undesirable.
46

0

5

10

15

20

25

30

35

40

45

0 5 10 15 20 25 30 35 40

Run Time, hr

200gpm

400 gpm

600gpm

800 gpm

Influent C

atio

n

C

o

ncentratio

n, meq/L

Figure 3-12: Influent cation concentration versus run time for Dow G-26 resin at a constant

resin volume equal to 1357 L at various feed water flow rates.

To overcome a very short run time when utilizing a non-continuous bed, the resin volume

is typically increased. By holding the time to regeneration constant the resin volume increases

with increasing cation concentration. This is illustrated in Figure 3-13 for a run time equal to

20hrs at various feed water flow rates. The result is an increase in plant size and capital costs.
47

0

5

10

15

20

25

30

35

40

45

0 10000 20000 30000 40000 50000 60000 70000 80000

Resin Volume, L

200gpm

400gpm

600gpm

800gpm

Influent C

atio

n

C

o

ncentratio

n, meq/L

Figure 3-13: Influent cation concentration versus required resin volume at a constant time to

exhaustion equal to 20 hrs for Dow G-26 resin.

3.4.4.3 Water-resin contacting methods

Water can contact the resin in four different configurations. Different water resin

contacting techniques include: 1) batch, 2) fixed bed, 3) fluidized bed (Wachinski, 2006), and 4)

moving bed (Drake, 2007).

Batch mode ion exchange works well for laboratory testing of resin when performing

isotherms, but is no longer employed beyond the lab (Wachinski, 2006). Batch ion exchange was

attempted at several locations in the PRB with zeolites, but was soon disbanded due to effluent

requirements not being met (Lamb, 2006).

Fixed bed ion exchange is ion exchange in which the resin bed does not move during the

service step. Fixed bed ion exchange can be either parallel flow or counter flow. Parallel flow

systems have a downward water flow during both service and regeneration steps. When utilizing
48

parallel flow one more step must be employed. Backwashing must be performed before

regeneration to remove particles that may get trapped in the resin during service, similar to a

sand filter which was discussed in pretreatment. Counter flow systems have a downward flow

service step and upward flow regeneration. Since this system regenerates similar to backwashing,

less water is required when compared to parallel flow (Dow, 2007).

Fluidized bed ion exchange is when the flow of water is upward (counter-current)

through the resin bed and the resin bed is allowed to expand (Wachinski, 2006). This

configuration does not require backwashing because particles cannot become entrained between

the media as a result of fluidization.

The final method for contacting water to resin is the moving bed configuration. In this

process the resin moves with the water to achieve a specified ion removal. A moving bed can

have flow regimes that exhibit elements of fluidization, slip flow and elutriation (Drake, 2007).

3.4.4.4 Waste Stream Management

Ion exchange systems can be configured with different waste management techniques.

There are four techniques used to minimize the volume of brine generated per service cycle.

Brine volume reduction techniques include: using more concentrated regenerant or a different

regenerant (Dow, 2007), reusing rinse water and/or regenerant (AWWA, 1990), and partial

regeneration (AWWA, 1990).

The concentration of the acid can be increased to the maximum recommended

concentration. Dow chemical company recommends an acid concentration between 0.205 to 2.18

eq/l (1% to 10% by weight) for sulfuric acid, and 1.12 to 2.28 eq/L (4 to 8% by weight) for

hydrochloric acid when using G-26 resin. By maximizing the recommended acid concentration

less regenerant waste is generated because more regenerative equivalents are present per volume
49

88%

90%

92%

94%

96%

98%

100%

0 5 10 15 20 25 30 35 40 45

Influent Cation concentration, meq/L

8% HCl

6% HCl

10% H2SO4

5% H2SO4

W

ater Reco

very

of regenerant, Figure 3-14. Using hydrochloric acid instead of sulfuric acid as the regenerant will

decrease the brine volume for the same reason.

Figure 3-14: Ion exchange water recovery rate as a function of influent cation concentration

regenerant type, and regenerant concentration as % by weight.

Another way to reduce the volume of brine generated is to reuse regenerant and or rinse

water instead of throwing it away. Rinse water is used to rinse the resin after regeneration. The

ultimate result is a very dilute acid solution that can be used for the next regeneration by simply

adding more acid to it. The result is a smaller volume of brine generated, Figure 3-15.

Regenerant can be reused if regeneration is performed in two steps, a roughing step and a

polishing step (AWWA, 1990). Regenerant reuse is not considered here due to the variability

surrounding it.
50

91%

92%

93%

94%

95%

96%

97%

98%

99%

100%

0 5 10 15 20 25 30 35 40 45

Influent Cation concentration, meq/L

8% HCl 0.25 new

8% HCl 0.5 new

8% HCl no reuse

W

ater Reco

very

Figure 3-15: Ion exchange water recovery rate as a function of influent cation concentration and

the fraction of water that is new per rinse or regeneration. For example, 0.25 new corresponds to

25% of the water used per regeneration or rinse not being used before.

Partial regeneration is the final way to minimize the amount of waste generated in an ion

exchange process. In this case the full capacity of the ion exchange resin is not restored during

regeneration. The result is effluent ion concentrations greater than what they would be if the

resin were fully regenerated. This will reduce the volume of brine generated, similar to blending

treated water with untreated water, but to what extent is unknown and the applicability of this

approach will most likely be governed by discharge requirements (AWWA, 1990). This method

is typically used when treating the entire water flow, i.e. no blending of treated and untreated

water (AWWA, 1990). The Drake Process is configured with partial regeneration (Drake, 2007).
51

3.4.5 PRB Ion Exchange Vendors

There are currently four ion exchange vendors in the PRB that are treating CBM water.

These vendors include: EMIT Water Discharge Technology, Rohm and Haas Company, Catalyx

Fluid Solutions Inc., and Drake Water Technology, LLC.

EMIT Water Discharge Technology has treated over 50,000,000 bbls of water in the PRB

(Beagle, 2006) using the Higgins Loop ion exchange configuration. The Higgins loop is

available in two sizes. The smaller one can treat 200-250 gpm of water and the large one can

treat up to 750 gpm depending on the influent ion concentration. This process generates less than

1% brine per volume of water treated and blended. The reason for such low brine generation is

the removal of brine from the loop when it reaches a certain concentration. Doug Beagle of

EMIT provided this information (Beagle, 2007).

Rohm and Haas operates a 400gpm fixed bed ion exchange system with sand filter

pretreatment in the PRB (Miers, 2007). The CBM water that they treat contains 1100 mg/L of

sodium and after treatment contains less than 5mg/L of sodium. The waste generated per volume

of water is approximately 2.5 - 3%. This corresponds to a water yield or recovery of 97 to 97.5%.

They are able to produce such little waste at such high cation concentrations because they reuse

fractions of rinse water. Jay Miers of Rohm and Haas provided this information (Miers, 2007).

Catalyx Fluid Solutions Inc. has a similar process as Rohm and Haas, but reuses

regenerant as well as rinse water. Catalyx is currently building an ion exchange water treatment

plant in the PRB.

Drake Water Technology utilizes the Drake Process. They have operated their system in

the PRB to verify the technology. This is a new technology that currently has five patents

pending. The skid mounted system can treat 250 gpm per unit. Higher flow rates can be
52

accommodated by running these systems in parallel. Brine generation is approximately 1% and

3% of the total flow for 15 and 55 meq/L cation removal, respectively. Ron Drake of Drake

Water Technology provided this information (Drake, 2007).

3.4.6 Costs

A non-continuous bed system will be used to determine the costs for ion exchange

because the continuous systems are not available without the vendor (EMIT or Drake). Based on

Dow recommendations a resin depth to vessel diameter ratio of 1:1 and service flow rate of 80

m/h will be used in the following analysis. Also, the run time will be assumed to be 20 hours

(Beardsley, 1994), and effluent cation and bicarbonate concentrations will be assumed to be zero.

All other anion effluent concentrations are assumed to be equal to their influent concentration.

Water recovery is determined by using Drake and Rohm and Haas water recovery data.

These data are illustrated in Figure 3-16. Water recovery is calculated with Equation 3.40

 100 - 0.05 (cation removal) r = ⋅ (3.40)

Where, cation removal is assumed to be equal to the influent cation concentration (meq/L).

y = 1 - 0.0005x

R

2

= 0.9798

96%

97%

98%

99%

100%

0 10 20 30 40 50 60

Cation Removal, meq/L

Water Recovery

Drake

Rohm & Haas

Figure 3-16: Ion exchange water Recovery as a function of cation removal. Data are from Drake

Water Technologies (2007) and Rohm and Haas Company (2007).
53

3.4.6.1 Capital Costs

Ion exchange system capital costs consist of: ion exchange vessels, resin, chemical

storage tanks, process water and brine holding tanks, chemical feed pumps and a structure to

house the system. These costs have been determined by Rubel (2003) through direct quotes from

manufacturers and equipment vendors.

The capital costs for an ion exchange treatment system are a function of the feed water

flow rate and the influent cation concentration when using a non-continuous bed system which

was the scope of Rubel (2003). Figure 3-17 was generated by adjusting Rubel (2003) data to

January 2007 dollars with ENR cost indexes.

y = 666.43x + 353856

R

2

= 0.9874

$100,000

$200,000

$300,000

$400,000

$500,000

$600,000

$700,000

$800,000

$900,000

$1,000,000

0 100 200 300 400 500 600 700 800 900

Resin Volume, ft

3

Capital Cost

Figure 3-17: Ion exchange capital costs as a function of resin volume for non-continuous ion

exchange. Adapted from Rubel (2003).

To determine the capital costs for a non-continuous ion exchange system Equation 3.36 is

used to size the system with respect to feed water flow rate and the influent cation concentration
54

by assuming a run time equal to 20hrs (Beardsley et al., 1994). Then the capital cost (CCIX) is

found using the correlation from Figure 3-17, Equation 3.41.

CCIX = 666.43 Vresin + 353856 (3.41)

3.4.6.2 Operating and Maintenance costs

Operating and maintenance costs for an ion exchange system include: Regeneration

chemicals, electricity, resin replacement, and labor. Operating and maintenance costs are

calculated with Equation 3.42.

OCIX = OCregen-IX + OCelec-IX + OCResin-IX + OClabor-IX(3.42)

Where, OCIX is the operating cost for an ion exchange system ($/bbl), OCregen-IX is the regenerant

cost, OCelec-IX is the electric cost, OCResin-IX is the resin replacement cost, and OClabor-IX is the cost

of labor.

The operating cost associated with regenerant depends on the feed water flow rate, cation

removal or influent cation concentration, and the regenerant purity or density. The cost of

regenerant in $/bbl is found by first determining the volume of regenerant required (L) per bbl of

feed water (bbl) with Equation 3.43.

f

regen

regen

Q

24hr/day

Regeneration

V

bbl water treated

V













⋅

















=

τ

(3.43)

Where, Qf is the feed water flow rate (bbl/day), and 
 is the time between regenerations (hr).

The Volume of concentrated/bulk acid required per barrel of feed water is then found

with Equation 3.44. This equation accounts for the dilution of the concentrated acid that is used

to make the regenerant solution.
55

acid

regen

regen

acid

EC

EC

bbl water treated

V

bbl water treated

V

⋅

















= (3.44)

Where, Vacid is the volume of concentrated/bulk acid required per barrel of feed water (L acid/bbl

water), ECregen is the equivalent concentration of the dilute acid/regenerant solution (2,870meq/L

for 10% hydrochloric acid (Dow, 2007)), and ECacid is the equivalent concentration of

concentrated/bulk acid (11,630meq/L for 36% hydrochloric acid (Dow, 2007)).

Finally, the operating cost associated with regenerant can be calculated. This is done with

Equation 3.45.

ρacid ⋅ ⋅











⋅











− = acid

acid

regen IX Cost

bbl water treated

V

ton

kg

OC 907.2 (3.45)

Where, Costacid is the unit cost of concentrated/bulk acid ($225/ton for 36% hydrochloric acid

(Miers, 2007)) and acid is the density of the concentrated/bulk acid (1.179kg/L for 36%

hydrochloric acid (Dow, 2007)).

Rubel (2003) estimated that the annual cost of electricity can vary between 0.05 and

0.033 kW-hr/1000 gal of feed water. The average electrical usage from Rubel (2003), 0.042 kWhr/1000gal or 0.0018kW-hr/bbl, is used to determine the annual cost of electricity with Equation

3.46.

OCelec-IX = (0.0018 kW-hr/bbl) Costelec (3.46)

Where, Costelec is the cost of electricity ($/kW-hr).

Resin replacement is approximately 20% of the resin volume per year, or a total resin life

of 5 years (Rubel, 2003; Beardsley et al., 1994). This cost was determined by Rubel (2003) to be

$0.05 to $0.20/ 1000gal depending on the plant size and method of shipping. The average will be
56

used to approximate ion exchange resin replacement costs and is assumed to be equal to

$0.12/1000 gal or $0.0053/bbl of feed water (Qf).

Rubel (2003) determined that a 750 ft

3

non-continuous ion exchange process with

automatic operations requires 465 man hours per year for routine operations and extra tasks. The

pay rate is assumed to be $50,000 per year for 2,000 man hours per year per employee. The

required number of man hours per year is assumed to be directly proportional to the volume of

resin required. The relationship between annual labor cost and resin volume is illustrated in

Figure 3-18. The cost of labor in $/bbl is calculated with Equation 3.47.

Q 365

15.5 V

OC

f

resin

Labor-IX

⋅

= (3.47)

Where, 15.5 Vresin is the annual cost of labor from Figure 3-18 ($/yr), Qf is the feed water flow

rate (bbl/day), and 365 is the number of days per year.

y = 15.5x

R

2

= 1

$0

$5,000

$10,000

$15,000

$20,000

$25,000

$30,000

0 500 1000 1500 2000

Resin Volume, ft

3

A

n

n

u

al L

a

b

o

r

C

ost

Figure 3-18: Annual labor cost as a function of resin volume for a non-continuous ion exchange

system, assuming annual labor cost is directly proportional to resin volume. Adapted from Rubel

(2003).
57

3.5 Electrodialysis

3.5.1 Overview

Electrodialysis (ED) is an electrochemical separation process in which ions are

transferred through charged membranes by means of an electrical potential difference

(Schoeman and Thompson, 1996). In this process ions are separated through the use of semipermeable ion-selective membranes (Metcalf & Eddy, 2003). Ion selective membranes allow

only anions or cations to cross them and can become even more selective when increasing or

decreasing the pore sizes in the membrane. For instance, membranes have been developed that

remove only monovalent ions from solution (Schoeman and Thompson, 1996). ED is illustrated

in Figure 3-19.

Figure 3-19: Electrodialysis process. Where, C is the cation selective membrane, A is the anion

selective membrane, Qt,i, and Qf,i are the influent and effluent flow rates for one flow channel, Cf

is the influent constituent concentrations, Ct is the effluent constituent concentrations, Cbr is the

effluent brine concentration, and Qbr,i is the brine flow rate for one concentrate flow channel.

Na

+

Na

+

Na

+

Na

+

Na

+

Na

+

Cl

-

Cl

-

Cl

-

Cl

-

Cl

-

Cl

-

A

C

C

A

Cl

-

Na

+

Cl

-

Na

+

+

-

Anode

Cathode

Qf,i

Cf

Qt,i

Ct

Qbr,i

Cbr

Qbr,i

Cbr
58

ED works by applying an electrical potential between two electrodes. This induces an

electrical current through the solution. The current then causes cations to migrate toward the

negative electrode and anions to migrate toward the positive electrode. As the ions migrate

toward their respective electrodes, the cation selective membrane allows cations to pass through

the cation selective membrane, and the anion selective membrane allows anions to pass through

the anion selective membrane. After the ions migrate through their respective membranes toward

the electrode of preference, they are stopped by the membrane of opposite charge (cations can

not migrate through an anion selective membrane). This generates a concentrate stream. The ions

can not diffuse back through the membrane that allowed them to enter the concentrate stream

because the current will not allow the ion to move against it.

ED has been used for the last fifty years to treat sea, brackish, and industrial waters

(Reahl, 2006). This process has also been used in the production of table salt, inorganic and

organic acids; and in the separation of amino acids from proteins, monovalent ions from

multivalent ions, and ionic compounds from uncharged molecules (Schoeman and Thompson,

1996).

3.5.2 Process

ED is typically performed in a stack (Schoeman and Thompson, 1996). A stack is a

module in which the membranes are oriented vertically and separated by flow spacers. Water

flows between the membranes which are separated by flow spacers. As the water flows through

the stack ions are transported through the membranes to a concentrate stream. The water is then

passed through another stack if further ion removal is required.

Several membrane stacks operated in parallel is typically referred to as a stage. The ion

removal per stage is approximately 50% of the total ion load to that stage (Schoeman and
59

Thompson, 1996). For example: the effluent TDS concentration for a three stage system with a

feed TDS equal to 1000 mg/L will be 500mg/L for stage 1, 250 mg/L for stage 2 and 125 mg/L

for stage 3. This process does not require blending for this reason. Blending is also not required

because the ion removal efficiency is adjustable by adjusting the voltage across the stack.

The efficiency of this process is determined by the concentration of ions in the

concentrate channels, the influent ion concentration, and the electrical potential, or voltage,

across the membrane stack. For these reasons electrodialysis is best characterized by current

utilization efficiency, Equation 3.48.

100%

NI

zFQ (C C )



f f,i t,i

×

−

= (3.48)

Where,  is the current utilization efficiency (%), z is the charge of the ion, F is Faraday’s

constant (96,485Amp-s/mol), Qf is the feed flow rate (L/s), Cf,i and Ct,i are the feed and treated

concentration for ion i, N is the number of cell pairs and I is the applied current (Amps).

To achieve the most economical operation of an ED system the current utilization

efficiency should be greater than 70% (EET, 2006). This is typically achieved by adjusting the

influent or effluent ion concentrations appropriately. The effluent concentration for a single stage

may be adjusted by increasing or decreasing the electrical potential (voltage) across the stack,

thus impacting the transport of ions through the membrane and increasing or decreasing the

effluent ion concentration. The influent ion concentration for ED is adjusted by either recycling a

portion of the concentrate through the stage in which it was produced or blending a portion of the

concentrate with the effluent from stage one as the water enters stage two. By increasing the

current utilization efficiency the electrical requirement will decrease (Schoeman and Thompson,

1996).
60

3.5.3 Electrodialysis Reversal (EDR)

EDR is the most commonly used ED configuration in water treatment (Reahl, 2006).

EDR removes constituents from the water in the same manner as regular ED except the polarity

of the stack is reversed every ten to fifteen minutes (Schoeman and Thompson, 1996). Polarity

reversal removes solids that may form on the membranes during normal operation.

The main benefit of this configuration is reduced chemical costs. When using EDR,

chemical pretreatment to consume alkalinity is not required because precipitates that may form

on the membranes are removed during polarity reversal. This results in reduced operating and

maintenance costs. Also, fewer chemicals are required to clean the membranes due to polarity

reversal cleaning the membranes to some extent (Schoeman and Thompson, 1996). For these

reasons EDR instead of regular ED should be used to treat CBM produced waters. Therefore, all

economic and performance analysis has been performed for EDR.

3.5.4 Chemical Rejection

The chemical rejection efficiency of an ED/EDR system is dependent on the number of

stages operated in series, the membrane type, the feed TDS of the water, and the current applied

(electrical potential across the stack). All of these variables have an impact on the effluent ion

concentrations and vary from facility to facility.

Schoeman and Thompson (1996) determined that chemical rejection efficiency for

ED/EDR facilities range between 40% and 60% per stage. The average of this range (50%) is

used to determine the chemical rejection efficiency for an ED/EDR system. The effluent ion

concentration from the system is calculated using Equation 3.49.

( ) f,i

N

Ct,i = 0.5 ⋅C (3.49)
61

Where, Ct,i is the effluent constituent concentration (mg/L), N is the number of stages required to

achieve the desired effluent concentration, and Cf,i is the feed constituent concentration (mg/L).

3.5.5 Water Recovery

Theoretically, water losses in an ED/EDR system are attributed to only the water

associated with ions (which form an ion’s hydraulic radius) as ions migrate through the

membrane (Schoeman and Thompson, 1996). Schoeman and Thompson (1996) related the

transport of water through the membrane and subsequently the water recovery (r) to the TDS

removal of an ED/EDR system. This relationship is described by Equation 3.50.

 100 -(5 10 ) (C -C ) f,TDS t,TDS

-4

r = × ⋅ (3.50)

Where, Cf,TDS is the feed TDS concentration (mg/L) and Ct,TDS is the treated water TDS

concentration (mg/L).

The theoretical water recovery was next compared to actual data collected by Leitner and

Associates (1997). Leitner and Associates (1997) data are illustrated in Figure 3-20. These data

do not provide a trend as predicted by the theoretical water recovery function. This difference

may be attributed to the systems from Leitner and Associates (1997) not operating to achieve the

highest possible water recovery and/or water being added to the concentrate flow channels to

reduce fouling potential and scale formation (Schoeman and Thompson, 1996). The averages and

standard deviations for Leitner and Associates (1997) data are 79% ± 8% for water recovery and

895 ± 660mg/L for TDS removal.

Based on data presented, the water recovery for an EDR system can be greater than 90%

if the system is optimized for water recovery, or below 80% for TDS removals equal to 900

mg/L as illustrated in Figure 3-20. The factor that has historically determined the water recovery

of an ED/EDR system has been the potential for precipitate formation (Schoeman and
62

Thompson, 1996). However, this is no longer as much of an issue due to recent developments in

membrane technology which allow ED/EDR systems to operate at constituent concentrations

well above saturation (Turek et al., 2006). Turek et al. (2006) achieved 93.1% water recovery

when operating an EDR at 364% calcium sulfate saturation.

y = -5E-06x + 1

R

2

= 1

60%

65%

70%

75%

80%

85%

90%

95%

100%

0 1,000 2,000 3,000 4,000 5,000 6,000 7,000 8,000

TDS Removed, mg/L

93%

94%

95%

96%

97%

98%

99%

Leitner and Associates, 1997 100%

Shoeman and Thompson, 1996

Sch

oem

a

n

a

n

d

T

h

o

m

pso

n

W

ater Reco

very

Leitner W

ater Reco

very

Figure 3-20: Water recovery versus TDS removed for 13 EDR systems (Leitner and Associates,

1997), and theoretical water recovery for an ED/EDR system (Schoeman and Thompson, 1996).

Due to recent developments in membrane technology and PRB CBM waters having

relatively low ( < 100mg/L) scale forming substances, the assumption has been made that EDR

systems can operate in the theoretical water recovery regime when treating PRB CBM waters.

Therefore, Equation 3.50 is used to calculate the water recovery for an EDR system.

3.5.6 Capital Costs

Capital costs for ED/EDR systems have been developed by Urano (1977), Larson and

Leitner (1979), Leitner and Associates (1997) and Lawrence et al. (1998). The key factors that
63

affect the capital cost of an EDR system are the required TDS removal and system capacity or

feed water flow rate (Schoeman and Thompson, 1996). Data from each source were adjusted to

January 2007 dollars using ENR cost indexes. These data are presented in Appendix A-ED.

Figures AED-1, AED-2, and AED-3 were developed with data from all of the above

sources to find EDR capital costs as a function of flow rate and TDS removal. EDR capital costs

(CCED) are calculated with Equations 3.51, 3.52, and 3.53 for TDS removals between 0 and less

than 2000mg/L, 2,000 and less than 7,000mg/L, and greater than or equal to 7,000mg/L,

respectively. Results for all of these functions are illustrated in Figure 3-21.

CC 0.0002 Q 90.36 Q 1000000 f

2

ED = ⋅f + ⋅ + 0 < TDSremoved < 2000mg/L (3.51)

CCED = 150.8 ⋅Qf +1000000 2000 TDSremoved <7000mg/L (3.52)

CCED = 355.9 ⋅Qf +1000000 7000mg/L TDSremoved (3.53)

$0

$2,000,000

$4,000,000

$6,000,000

$8,000,000

$10,000,000

0 5,000 10,000 15,000 20,000 25,000

C

a

pital C

ost, $

Feed Water Flow Rate, bbl/day

TDS >7,000

TDS 2000 to< 7000mg./L

TDS 0 to < 2000mg/L

Figure 3-21: EDR capital costs as a function of feed water flow rate and TDS removal as

predicted by Equations 3.51, 3.52, and 3.53.
64

3.5.7 Operating and Maintenance Costs

EDR operating and maintenance costs include: electricity, labor, chemicals, and

membrane and miscellaneous parts replacement. Methods used to determine these costs are

described below. The total operating and maintenance cost for an EDR system is calculated with

Equation 3.54.

OCED = OCelec-ED + OClabor-ED + OCchem-ED + OCmem-ED + OCparts-ED (3.54)

Where, OCED is the total operating and maintenance cost for EDR ($/bbl), OCelec-ED is the

electrical cost ($/bbl), OClabor-ED is the cost due to labor ($/bbl), OCchem-ED is the chemical cost

($/bbl), OCmem-ED is the membrane replacement cost ($/bbl), and OCparts-ED is miscellaneous parts

replacement cost ($/bbl).

The cost of electricity for an EDR system is dependent on the desired TDS removal and

the hydraulic pressure drop from stage to stage (Schoeman and Thompson, 1996). Leitner and

Associates (1997) and Lawrence et al. (1998) data have been used to develop a function that

relates electrical use to TDS removal. These data are illustrated in Figure 3-22. Equation 3.55 is

used to calculate the cost of electricity for an EDR system based on the desired TDS removal.

( ( ) ( )) elec

5 OCelec-ED = 0.0183 + Cf,TDS − Ct,TDS

⋅ 9×10 ⋅Cost

−

(3.55)

Where, OCelec-ED is the operating cost due to electricity ($/bbl), Cf,TDS is the feed TDS

concentration (mg/L), Ct,TDS is the treated water TDS concentration (mg/L), and Costelec is the

unit price of electricity ($/kW-hr).

The labor requirement for an EDR system has been developed from Leitner and

Associates (1997), and is assumed to be 1 full time employee per 7,732 bbl of feed water per

day. The cost for an employee (Costemployee) is assumed to be $50,000/yr. The cost of labor per

barrel of feed water (OClabor-ED) is calculated with Equation 3.56.
65

labor-ED Co employee

st

365 day

1 yr

day

bbl

7,732

1employee

OC ⋅

















⋅

























= (3.56)

Equation 3.56 shows that OCLabor-ED is independent of the flow rate. Therefore, the cost of labor

is considered fixed and equal to $0.0177/bbl.

y = 0.00009x + 0.0183

R

2

= 0.64419

0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

0 1000 2000 3000 4000 5000 6000 7000 8000

TDS Removed, mg/L

Leitner and Associates, 1997 Lawrence et al., 1998 Combined

Electrical Usa

ge, k

W-hr/b

bl

Figure 3-22: Electrical Usage with respect to TDS removal for 5 different EDR systems

operated at different TDS removals. Combined refers to the data set comprised of both data

sources (Leitner and Associates, 1997; Lawrence et al., 1998).

Chemical costs associated with an EDR system are typically attributed to periodic

cleaning of the membrane surfaces with dilute acids (Schoeman and Thompson, 1996). The

actual cost of chemicals at the Suffolk, VA EDR facility was $0.00042/bbl in 1994 (Schoeman

and Thompson, 1996). This corresponds to $0.00058/bbl in 2007 dollars. This value is used as

the chemical cost for EDR.
66

Membrane replacement is required every twelve to fifteen years if operated properly

(Reahl, 2006). Schoeman and Thompson (1996) have suggested that approximately 10% of the

membranes must be replaced annually, i.e. a ten year membrane life. Planners predicted the

membrane replacement cost at the Suffolk, VA EDR facility to be $0.06/1000gal or

$0.00252/bbl (Schoeman and Thompson, 1996). This value is used to estimate the cost due to

membrane replacement because it provides the most conservative estimate.

The final cost associated with EDR is miscellaneous parts replacement. This cost has

been estimated from Leitner and Associates (1997) data. Leitner and Associates (1997) collected

annual miscellaneous parts replacement and chemical data for different size EDR facilities. This

data was normalized with respect to flow to generate Figure 3-23. The average parts and

chemical cost from Figure 3-23 is $0.0021± 0.0013/bbl. The chemical cost from above

($0.00058/bbl) was subtracted from this value to yield the cost due to miscellaneous parts

replacement, $0.0015/bbl.

$0.0000

$0.0005

$0.0010

$0.0015

$0.0020

$0.0025

$0.0030

$0.0035

$0.0040

$0.0045

$0.0050

0 20,000 40,000 60,000 80,000 100,000 120,000

Feed Water Flow Rate, bbl/day

C

hemical a

n

d

P

arts C

ost, $/b

bl

Figure 3-23: Chemical costs and miscellaneous parts replacement cost normalized to feed water

flow rate. Adapted from Leitner and Associates (1997).
67

3.6 HEEPM

3.6.1 Process

High Efficiency Electro-Pressure Membranes (HEEPM) is a patent pending technology

marketed by EET Corporation. This technology combines electrodialysis and reverse osmosis.

After pretreatment, produced water is treated with electrodialysis then reverse osmosis.

Electrodialysis is used to reduce the ion load to reverse osmosis, and is also used to treat the

brine stream generated from reverse osmosis (EET, 2006). Reverse osmosis then further treats

the water to remove more constituents from solution. The process is outlined in Figure 3-24.

Electrodialysis Reverse Osmosis

Qf

Cf

Qt

Ct

Qbr-RO Cbr-RO

Qbr Cbr

Figure 3-24: Schematic for HEEPM water treatment system. Qbr-RO and Cbr-RO are the brine flow

and brine concentration being recycled from the RO unit to the ED unit.

This process has been used for water flows ranging from 1,000 gpd to 10 mgd, and is

adaptable to a wide range of influent TDS concentrations. A benefit to this process is higher

water recovery than is possible with RO alone, with a smaller system size than is required for ED

alone (Sferrazza, 2006). High recovery is achieved by recycling RO residuals through the

electrodialysis process. This concentrates the brine which results in less brine production. The

brine is removed from the system during electrodialysis.
68

Recycling RO residuals through the ED also helps to maintain a high ion concentration in

the ED. High ion concentrations in the ED influent water allows for higher current utilization

efficiency which is necessary for economic use of ED (Schoeman and Thompson, 1996). High

ion concentrations allow for less current to be absorbed by the water, thus allowing for less

energy being required to impart a current to the water.

HEEPM chemical rejection efficiencies are outlined in Table 3-7. These data were

supplied by EET Corporation, and are used to calculate the treated water chemistry (Ct).

Table 3-7: HEEPM chemical rejection efficiencies. Source: EET Corp.

Constituent Rejection Efficiency

TDS 98%

Chloride 98%

Sulfate 99%

Nitrate 90%

Phosphate 99%

Sodium 98%

Other 90%

HEEPM has the same water recovery as ED because the water and RO brine are treated

with ED. Therefore, water recovery for HEEPM and ED are related to feed TDS in the same

manner. Equation 3.57 is used to calculate the water recovery for HEEPM.

( ) f,TDS

-4

r =100 - 5×10 ⋅C (3.57)

3.6.2 HEEPM Costs

HEEPM capital and operating and maintenance costs were provided by EET Corporation

in two formats, estimates and plots. Two cost estimates were provided, one for a 15,000 bbl/day

system and one for a 45,000 bbl/day system. Each estimate contains information related to the

installed capital costs for each system in the PRB (see Appendix A-H Table AH-1). The

estimates also provide operating cost data such as membrane life and replacement cost, sulfuric
69

acid dose and cost, electrical use, and labor. In addition to the estimates, two plots were

provided. One relates operating and maintenance costs to feed water flow rate and raw water

TDS while the other relates capital costs to feed water flow rate and TDS. Both sources of

information were used to estimate the costs for HEEPM.

3.6.2.1 Capital Costs

Capital costs have been determined through the use of the provided plot and the two cost

estimates. The plot does not include costs such as, installation, freight, site preparation,

engineering, and storage tanks; but, the plot does provide a relationship which relates capital

costs to flow rate and TDS.

A strong relationship was not found to exist between TDS and capital cost for the TDS

concentrations of interest, 1,000 to 10,000 mg/L, in PRB CBM produced waters (see Appendix

A-H Figure AH-1). Therefore, average capital costs were found for this range of TDS values.

The averaged capital costs were then plotted with respect to feed water flow rate, Figure 3-25.

$10,000

$100,000

$1,000,000

$10,000,000

10 100 1,000 10,000 100,000

Feed Water Flow Rate (bbl/day)

C

a

pital C

ost, $

Estimate pts CC CC_adjusted

Figure 3-25: HEEPM capital costs with EET Corp. data from the plot relating capital costs, CC,

to feed flow rate; adjusted capital costs, CC_adjusted, takes into account installation, freight, site

preparation, engineering, and storage tanks; and estimate points (Estimate pts) from EET Corp.

capital cost estimates for two flow rates 15,000 and 45,000 bbl/day.
70

Data from the plot relating capital cost to flow rate were then adjusted to account for the

costs not included in the plot. This was done by finding common ratios between the actual

installed costs from the estimates and the data from the plot. In doing so, it was found that the

actual capital cost was 1.69 and 1.26 times greater for the 15,000 and 45,000 bbl/day systems,

respectively. Data from the plot were then adjusted by multiplying the capital cost from the plot

by: 1.69 for feed water flow rates greater than 2,381 and less than or equal to 23,810 bbl/day;

and 1.26 for flow rates greater than 23,810 bbl/day. The average of the ratios (1.48) was used for

flow rates less than 2,381 bbl/day. The adjusted capital costs are illustrated in Figure 3-25.

Functions were then found for the adjusted capital cost for different flow regimes by

plotting adjusted capital cost data versus feed water flow rate (Qf) and finding the best fit line for

each flow regime (see Appendix A-H Figures AH-2, AH-3, AH-4, and AH-5). HEEPM capital

costs (CCHEEPM ($)) are calculated with Equations 3.58, 3.59, 3.60 and 3.61.

CCHEEPM = 371.48⋅Qf + 75096 Qf 238 bbl/day (3.58)

CCHEEPM = 208.29⋅Qf +144883 238 < Qf 2381 (3.59)

CCHEEPM =160.99⋅Qf + 259932 2381 < Qf 23810 (3.60)

HEEPM 111 Qf CC = .47 ⋅ 23810 < Qf (3.61)

3.6.2.2 Operating and Maintenance Costs

HEEPM operating and maintenance costs include: electricity, labor, chemicals and

membrane replacement (EET, 2006). The total operating and maintenance cost per barrel of feed

water for HEEPM (OCHEEPM) is calculated with Equation 3.62.

OCHEEPM = OCelec−HEEPM + OClabor−HEEPM + OCchem−HEEPM + OCmem−HEEPM (3.62)

Where, OCelec-HEEPM is the electrical cost ($/bbl), OClabor-HEEPM is the labor cost ($/bbl), OCchem.-

HEEPM is the chemical cost ($/bbl), and OCmem-HEEPM is the membrane replacement cost ($/bbl).
71

The operating and maintenance costs associated with HEEPM were estimated from a

plot that relates total operating and maintenance cost to feed water flow rate and feed TDS

(Figure AH-6), and an estimate for treating a water with a feed TDS of 4400mg/L to 250 mg/L

provided by EET Corp (Table AH-2). With the use of both sources, the individual costs

represented in Equation 3.62 were uncoupled from Figure AH-6 and cost functions were found

for each component.

The electrical cost for HEEPM is directly related to the cost of electricity and the feed

TDS of the water (EET, 2006). The estimated electrical use to treat a water with 4400 mg/L feed

TDS concentration is 1.16 kW-hr/bbl (EET, 2006). This electrical usage was used in conjunction

with Figure AH-6 data to develop a relationship between feed TDS and electrical usage (See

Appendix A-H). This relationship is illustrated in Figure 3-26. The electrical cost for HEEPM is

calculated with Equation 3.63.

y = 0.0002x + 0.4373

R

2

= 0.9975

0

1

2

3

4

5

6

0 5,000 10,000 15,000 20,000 25,000 30,000 35,000 40,000

TDS, mg/L

Electrical Usa

ge, k

W-hr/b

bl

Figure 3-26: HEEPM electrical usage as a function of feed TDS.
72

( ) elec HEEPM f,TDS C elec OC = 0.0002⋅C + 0.4373 ⋅ ost

−

(3.63)

Where, Cf,TDS is the feed TDS concentration (mg/L) and Costelec is the unit cost of electricity

($/kW-hr).

The operating and maintenance cost attributed to labor for HEEPM is assumed to be

equal to $0.0151/bbl of feed water (EET, 2006). This cost assumes an employee cost of

$50,000/employee-yr.

HEEPM requires acid addition for the same reason as RO (to consume alkalinity which

prevents membrane fouling in the RO unit). The amount of acid required from the estimate is

0.00071 tons of 98% pure sulfuric acid per barrel of feed water for a water containing

approximately 5078 mg/L of alkalinity as calcium carbonate (see Appendix A-H for an

explanation of this feed water alkalinity). This is approximately equal to the acid required from

Figure 3-10. The sulfuric acid requirement is assumed to be directly proportional to the feed

TDS, Figure 3-27 (which is identical to Figure 3-10 with the exception of the x-axis being TDS

concentration for Figure 3-27). Chemical cost is calculated with Equation 3.64.

f,TDS acid

7 OCchem HEEPM = 2×10 ⋅C ⋅Cost−

−

(3.64)

HEEPM membrane replacement cost is assumed to be $0.0126/bbl of feed water (EET,

2006). This cost assumes a 3 and 5 year membrane life for RO and ED membranes, respectively.
73

y = 2E-07x

R

2

= 1

0.0000

0.0005

0.0010

0.0015

0.0020

0 2,000 4,000 6,000 8,000 10,000

TDS, mg/L

T

o

ns S

ulfu

ric

A

cid

p

e

r

b

bl w

ate

r

Figure 3-27: HEEPM 98% pure sulfuric acid requirement as a function of feed TDS assuming

acid use is directly proportional to feed TDS.
74

4. BRINE MANGEMENT

4.1 Overview

All of the selected water treatment technologies generate a waste stream as a result of

removing constituents from water. This waste stream is generally termed brine. The brine

contains all of the constituents that were removed from the water and as a result must be

managed in a manner that will not release these constituents into the environment.

Two brine management techniques are currently used in the PRB. These include deep

well injection and brine evaporation ponds (Lamb, 2006). Other techniques that may be used to

manage brine in the PRB include: evaporation crystallization, bipolar electrodialysis, and

membrane distillation.

4.2 Deep Well Injection

4.2.1 Overview

Deep well injection is a brine management technique that involves injecting brine into a

geologic formation in the subsurface. The formation in which the brine is injected must be

beneath the lowermost underground source of drinking water, and must not allow the brine to

migrate to other areas of the subsurface which may result in aquifer contamination (Mickley,

2001; ASCE/AWWA, 1996). Wells used for this purpose are termed Class I injection wells and

are typically 1,000 to 8,000 feet deep (Mickley, 2001).

This brine management technique has been used in the PRB for several years with much

success (Lamb, 2006). The primary issue with deep well injection is that there is a set capacity
75

for how much brine a formation can store (Beagle, 2007). When all of the wells that are used for

deep well injection have filled the capacity of the formation in which they inject brine either: a)

more wells will have to be developed, or b) another brine disposal technique will have to be

used.

Class I injection wells are regulated by the WDEQ in Wyoming under Chapter 13 of the

Wyoming water quality rules and regulations, and are nationally regulated by the US EPA under

40 CFR Parts 144 and 146. Currently there are 50 class I injection wells in Wyoming that are

permitted for non-hazardous wastes (Passehl, 2007). Wyoming does not have any class I wells

that are regulated for hazardous wastes. Of the 50 class I wells in Wyoming, less than five are

used for CBM water treatment residuals. One of these wells is used exclusively by EMIT

Discharge Technologies (Beagle, 2007).

Beagle (2007) has supplied information for the well that is used exclusively by EMIT

Discharge Technologies. This well is 9,178 to 9,282 ft deep in the Minnelusa formation of

northern Wyoming and permitted for 5,040 bbl/day of brine at 1,700 psi maximum injection

pressure. To date EMIT has injected approximately 900,000 barrels of brine, and the actual

capacity of the formation is unknown (Beagle 2007). Cost information for this well was not

available due to the proprietary nature of the costs. However, this well’s depth and flow rating

will be used to develop costs for deep well injection.

4.2.2 Costs

There are several costs associated with the construction of a Class I injection well that

can be used to dispose of brine. These costs include: subsurface exploration for a formation that

is capable of isolating and storing the brine, permitting, materials, and well installation (Mickley,

2001). Mickley (2001) developed a cost model that predicts the capital costs associated with the
76

installation of a Class I injection well (Equation 4.1). This model predicted the total installed cost

of a 9,300 ft deep Class I injection well with a 16 inch injection tube diameter (capable of

injecting 4.51 to 9.02 mgd) to be $9,058,600, not including permitting.

CCW 288 145.9 (Tube Diameter) 0.754 (Depth) ell = − + ⋅ + ⋅ (4.1)

Where, CCWell is the capital cost for a class I injection well ($1,000), Tube Diameter is the

injection tube diameter (in), and Depth is the well depth (ft).

In addition to capital costs there are also operating costs that must be considered when

utilizing a Class I injection well. These costs include: pumping, chemicals and labor. The most

significant of these costs is pumping power which is estimated to be $50,000/yr for a 150 gpm

(5142 bbl/day) pump operating at 3,150 psig with a 350 hp motor (Mickley, 2001).

Chemicals are required to prevent corrosion of the pump and injection tube. Mickley

(2001) estimated that these costs are approximately $7,000/yr for a 150 gpm injection well.

Labor costs are neglected because an injection well requires labor primarily for monitoring

activities and periodic repairs.

Based on Mickley (2001) results, the cost to dispose of a barrel of brine if the company

owns the well will be $0.26/bbl (not including trucking or permitting) assuming: the well

capacity is 1,646 million gallons, capital cost is $9,058,600, flow rate is 150 gpm (5143 bbl/day)

for electric and chemical costs, and 21 year straight line depreciation. Currently, the cost to

dispose of a barrel of brine via deep well injection in the PRB is $5.00/bbl including trucking the

brine to the well if the well is not owned by the company disposing of the brine (Miers, 2007).
77

4.3 Evaporation Ponds

4.3.1 Overview

Evaporation ponds utilize solar energy to evaporate water from brine. When designed

appropriately the end result is dried solids that may be landfilled depending on the characteristics

of the solids. Evaporation ponds are typically used in dry climates with high evaporation rates to

ensure that water will evaporate from the brine (Mickley, 2001). Evaporation ponds are currently

being used at Petro-Canada’s Wild Turkey RO facility to manage brine (Lamb, 2006).

This technique has several issues related to its use. Issues include: Large evaporative area

requirement for large flows, freezing in the winter, potential underlying aquifer contamination

from flooding or damaged pond liners (Mickley, 2001), microbial and/or algae growth (Lamb,

2006), and solids disposal. There are some advantages, however, to using an evaporation pond.

These include: relatively easy to construct, low maintenance costs, no chemicals required, and no

mechanical equipment required except for a pump to get the brine to the pond and possibly a

water atomizer/evaporator to enhance evaporation (Mickley, 2001).

4.3.2 Parameters

Important parameters for evaporation ponds include the pond liner thickness, dike height,

evaporative area, evaporation rate and pond depth. The thickness of the pond liner is important to

ensure the liner will not become damaged and contaminate underlying groundwater. Mickley

(2001) reported that liner thicknesses can range from 20 to 120 mils and that the thickness may

be regulated by law. A liner thickness equal to 60 mils is assumed to be adequate (Mickley,

2001).
78

Dikes are necessary around the perimeter of a pond to ensure the brine stays in the pond

during times of strong winds and precipitation. The dike height is assumed to be 8 feet above the

pond surface (Mickley, 2001).

Adequate evaporative area is also necessary to provide surface area for evaporation to

occur. The required evaporative area is dependent on the local evaporation rate which is 2.51

ft/yr and 5.80 ft/yr when utilizing natural or forced evaporation in the PRB, respectively (Urruty,

2007). Forced evaporation is the use of an atomizer/evaporator to enhance evaporation. The

evaporation rate for water saturated with sodium chloride salt has been found to be 70 percent of

the evaporation rate for freshwater (USDI, 1971). Therefore the evaporation rate for an

evaporation pond utilizing forced evaporation in the PRB is assumed to be 4.06 ft/yr. The

required evaporative area is calculated with Equation 4.2.

( ) ( )

ER

Q

EA 1.2 2049.5

br = ⋅ ⋅ (4.2)

Where EA is the evaporation area (ft

2

), 1.2 is a 20% safety factor, 2049.5 is a conversion factor

from bbl yr ft

-1

day

-1

to ft

2

, Qbr is the brine flow rate (bbl/day), and ER is the evaporation rate

(ft/yr).

The depth of the pond is also important. Pond depth provides storage for solids that

precipitate out of the brine as the water evaporates. The rate of solids accumulation can range

from 0.005 to above 0.02 ft/yr for brine TDS concentrations ranging from 15,000 to above

50,000 mg/L. A total pond depth equal to 1.5 feet is assumed to be adequate for solids storage

(Mickley, 2001).
79

4.3.3 Costs

Seth Urruty of Petro-Canada has supplied cost data for the Wild Turkey evaporation

ponds in the PRB. These data have been used to approximate capital, and operating and

maintenance costs for evaporation ponds.

Petro-Canada currently has four evaporation ponds with a total evaporation area of

818,700 ft

2

in the PRB (Urruty, 2007). These ponds are used to evaporate brine generated from a

RO facility that achieves 92% to 96% water recovery. The total design flow rate for the

evaporation ponds is 80 gpm (Urruty, 2007).

Costs associated with constructing an evaporation pond include: land, earthwork, lining,

and other miscellaneous costs (Mickley, 2001). The total capital cost to construct 818,700 ft

2

of

evaporation pond in the PRB was $1,500,000 ± $150,000 not including the cost of land (Urruty,

2007). Urruty (2007) stated that it is not required to purchase land for an evaporation pond in the

PRB, but that purchasing land may shield the producer from future liability issues. The cost of

land for this purpose in the PRB is confidential (Urruty, 2007). The unit cost to construct an

evaporation pond in the PRB is then $1.83/ft

2

. The capital cost associated with evaporation pond

brine management is calculated with Equation 4.3 (assuming capital cost is directly proportional

to evaporative area).

CCpond =1.83⋅EA (4.3)

Where, CCpond is the evaporation pond capital cost ($), 1.83 is the unit capital cost ($/ft

2

), and

EA is the evaporative area (ft

2

).

Operating and maintenance costs for the Wild Turkey ponds are approximately

$48,000/yr when utilizing forced evaporation (Urruty, 2007). The total operating and

maintenance cost per barrel of brine managed via forced evaporation is $0.081/bbl assuming the
80

evaporation rate is 4.06 ft/yr for the entire 818,700 ft

2

evaporative area. This is the assumed

operating and maintenance cost for evaporation pond brine management in the PRB.

4.4 Evaporation Crystallization

4.4.1 Overview

Evaporation is a process which is used to desalinate seawater around the world and

crystallization is used to generate solids. These are separate processes, but are generally used

together when managing brine. Evaporation is referred to as a process in which energy is applied

to water causing it to evaporate. Evaporation and crystallization can occur when solar energy is

applied as is the case with evaporation ponds. Evaporation crystallization can also occur as a

result of applying energy to water in a controlled environment, or as a result of atomizing the

water (spray evaporation).

When using evaporation crystallization to manage brine, evaporators concentrate the

brine from 1 - 3.5% total dissolved solids to 10-20% total solids by weight (ASCE/AWWA,

1996). The concentrated brine is then passed through a crystallizer that can concentrate and

solidify the brine to 50% total solids by weight. The brine solids extracted from the crystallizer

may then be dewatered further. The solids are then landfilled or sold as a product.

4.4.2 Process

There are several types of evaporators that can be used to concentrate brine. Evaporators

include: multi-effect distiller (MED), multi-stage flash (MSF), vertical tube falling film vapor

compression, rapid spray, freeze thaw, and solar.

Mickley (2001) and ASCE/AWWA (1996) recommend the use of a vertical tube falling

film vapor compression evaporator to concentrate brine. This type of evaporator is recommended
81

because it requires less energy and has lower capital costs than the other types of evaporators.

This type of evaporator is illustrated in Figure 4-1.

Compressor

Brine Recirculation Pump

Concentrated Brine to

Crystallizer

Vapor

Brine

Brine In

Distillate to heat

exchanger

D

e

a

erator

Tubes

Brine

Gases

Distillate

Out

Brin

e

H

e

at

E

xc

ha

n

ger

E

v

a

p

orator

Figure 4-1: Vertical tube falling film vapor compression evaporator. Adapted from

ASCE/AWWA, 1996.

Vertical tube falling film vapor compression evaporators utilize a vapor compression

cycle to minimize heat inputs. This type of evaporator requires heat energy to bring the brine to a

boil at start-up. After the system is up to operating temperature the external heat source is no

longer needed. Vapor that is generated as a result of evaporation is compressed to maintain the

temperature within the evaporator. This type of evaporator is 28 times more efficient than a

steam powered evaporator (ASCE/AWWA, 1996).

Vertical tube falling film vapor compression evaporators should operate with a feed TDS

concentration between 10,000 and 35,000 mg/L to prevent scaling of components. First, feed

brine is pumped through a heat exchanger, which has hot distillate circulating through it, to
82

preheat the brine. Brine then flows through a deaerator that vents non-condensable gases that

may interfere with heat transfer, carbon dioxide that may cause scale formation, and oxygen to

prevent corrosion (ASCE/AWWA, 1996). Brine then falls through the vertical tubes which are

contacted by compressed vapor on the outside of the tubes. Most of the water evaporates (91.8%

total when concentrating from 2 to 20% total solids) and leaves the system as distillate (TDS =

10 to 50 mg/L). Remaining vapor is then compressed to increase the temperature of the vapor by

10 to 15oF above the boiling temperature in the evaporator (ASCE/AWWA, 1996). The

compressed vapor is then released back into the evaporator to evaporate more water from the

brine and maintain the evaporator’s operating temperature. The brine is recirculated, at 20 to 40

times the feed flow rate, through the evaporator to maintain the temperature in the evaporator

and further concentrate the brine. The brine leaves the evaporator with a TDS concentration

around 200,000 mg/L (ASCE/AWWA, 1996).

Once the brine has a TDS concentration around 200,000 mg/L it leaves the

evaporator/concentrator and enters the crystallizer. The crystallizer uses compressed steam to

further concentrate the brine to a TDS concentration around 500,000 mg/L. At this concentration

the brine slurry begins to crystallize. The slurry is then extracted from the crystallizer and

mechanically dewatered with a centrifuge, belt filter, filter press, hydrocyclone, or evaporation

pond to achieve final brine solids concentrations of 80-90%, 40-50%, 40-60%, 50-70%,

respectively (ASCE/AWWA, 1996).

The use of an evaporator is required when Cbr-TDS is below 200,000 mg/L as illustrated in

Figure 4-2a. If Cbr-TDS is greater than or equal to 200,000 mg/L an evaporator is not required to

concentrate the brine for the crystallizer as illustrated in Figure 4-2b. A mass balance on Figures
83

4-2a and 4-2b allows for the crystallizer design flow rate (Qbr-1) to be determined (Equation 4.4

and Equation 4.5).

br TDS-1

br br TDS

br-1

C

Q C

Q

−

−

⋅

= Cbr-TDS < 200,000 mg/L (4.4)

Qbr-1 = Qbr Cbr-TDS  200,000 mg/L (4.5)

Where, Qbr is the brine flow rate leaving the demineralization technology (bbl/day), Qbr-1 is the

concentrated brine flow rate entering the crystallizer (bbl/day), Qbr-2 is the brine flow rate leaving

the crystallizer (bbl/day) , Cbr-TDS is the brine TDS concentration leaving the demineralization

technology (mg/L), Cbr-TDS-1 is the brine TDS concentration leaving the evaporator (200,000

mg/L), and Cbr-TDS-2 is the brine TDS concentration leaving the crystallizer (500,000 mg/L).

Qbr

Cbr-TDS

Demineralization

Technology

Evaporator Crystallizer

Qbr-1

Cbr-TDS-1

Qbr-2

Cbr-TDS-2

Qbr

Cbr-TDS

Demineralization

Technology

Crystallizer

Qbr-2

Cbr-TDS-2

a)

b)

Figure 4-2: Evaporation crystallization process for brine leaving the demineralization

technology with a TDS concentration a) less than 200,000 mg/L and b) greater than or equal to

200,000 mg/L. Variables are defined below Equation 4.5.

4.4.3 Costs

Evaporation crystallization costs have been adapted from Mickley (2001) and

ASCE/AWWA (1996). The costs for each unit have been assessed in detail and adjusted with

ENR cost indexes to January 2007 dollars. Costs include: evaporator capital, evaporator
84

operating and maintenance, crystallizer capital, and crystallizer operating and maintenance. The

overall cost for evaporation crystallization is calculated with Equation 4.6.

t Q 365

CC CC

C OC OC

br

e c

e/c e c

⋅ ⋅

+

= + + (4.6)

Where Ce/c is the cost to manage a barrel of brine with evaporation crystallization ($/bbl), OCe is

the evaporator operating cost ($/bbl), OCc is the crystallizer operating cost ($/bbl), CCe is the

evaporator capital cost ($), CCc is the crystallizer capital cost ($),t is the service life of the

facility (yrs), Qbr is the brine flow rate (bbl/day), and 365 is the number of days in a year.

4.4.3.1 Evaporator

Capital costs for vertical tube falling film vapor compression evaporators are illustrated

in Figure 4-3. The trend line for Mickley (2001) data is used to calculate evaporator capital cost

because the combined trend line and ASCE/AWWA (1996) trend line both imply that the

minimum cost for an evaporator is $2,000,000. This is because ASCE/AWWA (1996) data do

not include capital costs for brine flow rates below 100 gpm. Mickley (2001) data does include

capital costs for brine flow rates as low as 25 gpm. Therefore, Mickley (2001) data is used to

calculate evaporator capital cost. Equation 4.7 is used to calculate the capital cost for a vertical

tube falling film vapor compression evaporator/concentrator.

CCe = 330 Qbr + 738,708 (4.7)

Where CCe is the evaporator capital cost ($) and Qbr is the brine flow rate entering the evaporator

(bbl/day).

Operating and maintenance costs associated with an evaporator include: energy to heat

the brine and compress the vapor, energy for pumping fluids, and labor. Mickley (2001) found

that the total energy requirement for an evaporator is 3.8 kW-hr/bbl of brine. ASCE/AWWA

(1996) found that the overall energy requirement can range from 4.62 to 7.35 kW-hr/bbl. The
85

average of these values (5.26 kW-hr/bbl) is assumed to be the energy required to concentrate

brine from a TDS concentration of 20,000 to 200,000 mg/L. This is approximately equal to

0.0367 kW-hr/L of distillate produced. The energy required to evaporate pure water without an

evaporator would be approximately 0.725 kW-hr/L of distillate (See Appendix B). Therefore,

this type of evaporator is much more efficient than simply applying heat energy to evaporate

water.

Combined: y = 312.33x + 2E+06

R

2

= 0.9454

ASCE/AWWA, 1996: y = 443.01x + 2E+06

R

2

= 0.9883

Mickley, 2001: y = 330x + 738708

R

2

= 0.9862

$0

$5,000,000

$10,000,000

$15,000,000

$20,000,000

$25,000,000

$30,000,000

0 10,000 20,000 30,000 40,000 50,000 60,000 70,000 80,000

Brine Flow Rate, bbl/day

C

a

pital C

ost

Mickley, 2001

ASCE/AWWA, 1996

Combined

ASCE/AWWA, 1996

Mickley, 2001

Figure 4-3: Capital cost with respect to influent brine flow rate for a vertical tube falling film

vapor compression evaporator. Adapted from Mickley (2001) and ASCE/AWWA (1996).

Combined is the trend line that is comprised of data from both sources.

Figure 4-4 illustrates the energy requirement with respect to the evaporator feed TDS.

This figure was produced by assuming the energy required to concentrate brine to a TDS

concentration equal to 200,000 mg/L is inversely proportional to the feed TDS. Equation 4.8 is

used to calculate the energy requirement to concentrate brine to 20% solids.

( ) br TDS

-5 Ee

5.9 2.95 10 C − = − × ⋅ (4.8)
86

Where, Ee is the energy required to concentrate brine to 20% solids (TDS = 200,000 mg/L) and

Cbr-TDS is the influent brine TDS concentration (mg/L).

y = -0.0000295x + 5.8889

0

2

4

6

0 50,000 100,000 150,000 200,000

Evaporator Feed TDS, mg/L

E

nerg

y

Req

uired, k

W-hr/b

bl

Figure 4-4: Energy required to concentrate brine to 20% solids. Assuming the energy required is

inversely proportional to the feed TDS.

The labor requirement is assumed to be 2 to 4 hrs per eight hour shift (Mickley, 2001).

This is equivalent to 520 to 1,040 hrs per year (assuming 5 shifts per week), or $5,200 to $10,400

per year (assuming an employee cost (Cemployee) equal to $50,000 per 2000 hrs per year or

$25/hr). The overall operating and maintenance cost for evaporation/concentration is calculated

with Equation 4.9.

br

employee

e e elec Q

4 C

OC E C

⋅

= ⋅ + (4.9)

Where, OCe is the operating and maintenance cost for an evaporator ($/bbl), Ee is the electrical

usage (kW-hr/bbl), Celec is the unit price of electricity ($/kW-hr), 4 is the labor requirement per

day (hr), Cemployee is the employee pay rate ($/hr), and Qbr is the brine flow rate (bbl/day).
87

4.4.3.2 Crystallizer

Crystallizer capital costs have been adapted from Mickley (2001) and ASCE/AWWA

(1996), and are illustrated in Figure 4-5. The capital cost for a crystallizer is calculated with

Equation 4.10.

0.5521

c 67 Qbr-1 CC = ,582⋅ (4.10)

Where, CCc is the capital cost for a crystallizer ($) and Qbr-1 is the brine flow rate entering the

crystallizer (bbl/day).

y = 67582x

0.5521

R

2

= 0.9771

$0

$500,000

$1,000,000

$1,500,000

$2,000,000

$2,500,000

$3,000,000

$3,500,000

$4,000,000

$4,500,000

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800

Concentrate Flow Rate (Qbr-1) , bbl/day

C

a

pital C

ost

Mickley, 2001

ASCE/AWWA, 1996

Combined

Figure 4-5: Crystallizer capital costs with respect to concentrate flow rate (Qbr-1). Adapted from

Mickley (2001) and ASCE/AWWA (1996). Combined is the trend line that is comprised of data

from both sources.

Operating and maintenance costs associated with crystallizers are similar to those for

evaporators/concentrators. The energy required to crystallize brine from a TDS concentration of

200,000 to 500,000 mg/L was found to range from 8.4 to 10.5 kW-hr/bbl (Mickley, 2001;

ASCE/AWWA, 1996). The average (9.45 kW-hr/bbl) is assumed to be the energy requirement.
88

This is equal to 0.0992 kW-hr/L of distillate produced. The energy required to evaporate pure

water without this type of crystallizer would be approximately 0.63 kW-hr/L of distillate

assuming the only energy required is for vaporization (heat of vaporization = 2270 kJ/L).

Therefore, this type of crystallizer is much more efficient than simply applying heat energy to

evaporate water.

Figure 4-6 illustrates the energy requirement with respect to the crystallizer feed TDS.

This figure was produced by assuming the energy required to concentrate brine to a TDS

concentration equal to 500,000 mg/L is directly proportional to the crystallizer feed TDS (CbrTDS-1) . Equation 4.11 is used to calculate the energy required to crystallize/concentrate brine

from Cbr-TDS-1 to 50% solids.

c 0 Cbr TDS-1 E 15.75 .00003 − = − ⋅ (4.11)

Where, Ec is the crystallizer energy requirement (kW-hr/bbl) and Cbr-TDS-1 is equal to Cbr-TDS for

Cbr-TDS  200,000 mg/L.

The labor requirement for a crystallizer is assumed to be the same as that for an

evaporator. The total operating and maintenance cost for a crystallizer is calculated with

Equation 4.12.

br 2

employee

e c elec Q

4 C

OC E C

⋅

= ⋅ + (4.12)

Where Ec is the energy usage (kW-hr/bbl), Qbr 2 is the concentrated brine flow rate entering the

crystallizer (bbl/day), Celec is the unit price of electricity ($/kW-hr), 4 is the hourly labor

requirement per day (hr), Cemployee is the hourly employee pay rate ($/hr).
89

y = -3E-05x + 15.75

0

2

4

6

8

10

200,000 250,000 300,000 350,000 400,000 450,000 500,000 550,000

Crystallizer Feed TDS, mg/L

E

nerg

y

Req

uired, k

W-hr/b

bl

Figure 4-6: Energy required to crystallize/concentrate brine from Cbr-TDS-1 to 50% solids.

Assuming the energy required is inversely proportional to the feed TDS.

4.5 Other Options

4.5.1 Bipolar Electrodialysis

Bipolar electrodialysis (BED) is a technology that has been used to manage ion exchange

brine streams and is also used to produce acids and bases commercially (Schoeman and

Thompson, 1996). This technology functions similarly to an ED unit used to demineralize water

with a modification. This modification is a bipolar membrane which is comprised of both an

anion layer that faces the anode and cation layer that faces the cathode. The bipolar membrane

allows water molecules to enter it and as a result of the membrane orientation and electrical field

generated by the cathode and anode the bipolar membrane splits water molecules into hydrogen

and hydroxyl ions (Eurodia, 2007). These ions migrate toward their respective ends of the ED

stack and form a basic and an acidic solution. Bipolar electrodialysis is illustrated in Figure 4-7.
90

C

A

A

C

+

-

Anode

Cathode

H2O

Bipolar Membrane

H2O

H+

OHHCl

NaOH

NaCl

NaCl

Na

+ Cl

-

Cl N -a

+

Figure 4-7: Bipolar electrodialysis (BED). NaCl represents the brine feed, HCl is the acid

produced, NaOH is the base produced, A is the anion selective membrane, C is the cation

selective membrane, H

+

and OH

-

are the ions generated from H2O by the bipolar membrane.

Adapted from Wisniewski et al. (2004).

The feed solution for BED is the brine generated from ion exchange or another

demineralization technology and water. Ion exchange brine management may be the most

economically viable use for BED though. This is due to ion exchange brines having relatively

pure concentrations of sodium chloride or sodium sulfate when treating PRB CBM waters.

The end result of treating PRB CBM ion exchange brines with BED is an acid

(hydrochloric acid or sulfuric acid) and a base (sodium hydroxide). If managed properly BED

generates a product that is necessary to treat these waters via ion exchange, acid. As a result, the
91

once waste brine has become a useful consumable, acid that can be used to regenerate cation

exchange resins. In addition to acid generation, the base may be sold as a product.

This technology has been used extensively in Europe to produce acids and bases from ion

exchange brines, but its use has been limited in the United States. Eurodia Industries currently

markets BED in the United States (Eurodia, 2007). Cost information was not available, but

should be further investigated by interested parties.

4.5.2 Membrane Distillation

Membrane distillation is a process that may be used to concentrate brine similarly to an

evaporator. The process involves the transport of water vapor through the pores of hydrophobic

membranes due to a vapor pressure driving force provided by temperature and/or solute

concentration differences across a membrane (Martinez-Diez and Florido-Diaz, 2000). This

process is illustrated in Figure 4-8.

Figure 4-8 shows the higher temperature on the high TDS side of the membrane causing

water to enter the gas phase and leaving solids in the liquid phase. The vapor then crosses the

membrane to the lower temperature side and becomes liquid again as a result of the gas not being

able to hold as much water at the lower temperature.

High

TDS

Low

TDS

M

e

m

bra

n

e

Vapor

50oC 30oC

Figure 4-8: Membrane distillation when using a temperature difference as the driving force to

produce a lower TDS product.
92

This process has been used to desalinate drinking water from brines (Martinez-Diez and

Florido-Diaz, 2000; Drioli et al., 1999; Godino et al., 1996), but has not been used for the sole

purpose of concentrating brines to date. When used to produce drinking water the brine is heated

to a temperature between 30 and 50 oC. This is a much lower temperature than required for an

evaporator which requires the brine to boil at temperatures in excess of 100 oC. Therefore, this

process may offer some advantages and be more economical than an evaporator.
93

5. TOOLBOX

5.1 Overview

The Coalbed Methane Demineralization Treatment Cost (CBM-DTC) Toolbox has been

developed to predict the costs associated with treating PRB CBM waters based on influent water

characteristics, effluent water requirements, and the abilities of selected technologies. The costs

predicted by the toolbox should not be considered to be the exact costs, but approximations.

The toolbox is constructed in Visual Basic with the user interface being Excel based. This

configuration keeps the user interface relatively small with respect to the size of the logic

required to perform required calculations, thus minimizing spreadsheet calculations and the

probability of user created errors. The presence of visual basic macros are identified by a series

of buttons located though out the spreadsheet (Table 5-1).

Table 5-1: Technology and brine management user interface with buttons shown under “Click to

Select a Technology” and “Click to Select Brine Management”. Numbers shown are for example

only.

Technology Output: Brine Output:

Annual Per bbl

Electric $468,278 $0.013

Labor $426,888 $0.012

Chemical $2,228,356 $0.061

Membrane/Parts $791,522 $0.022

Resin $0 $0.000

Iron Pretreatment $251,136 $0.007

Total O&M $4,166,181 $0.114

Total Capital $5,347,150 $0.015 Total, $/bbl of brine $5.00

Total w/o Brine $0.129 Brine Cost/bbl of Qi $0.171

Click to Select a Technology Click to Select Brine Management

Total Treatment Cost

Reverse Osmosis

Deep Well

Injection

$0.300

IX HCl IX H2SO4 EDR HEEPM RO

Deep Well

Injection

Evap/Cryst.

Evap.

Pond
94

To use the toolbox the user must answer a series of questions by entering an answer in

the adjacent cell. Then to run a specific macro the user must hit enter then click the mouse on the

necessary button. The results will then be displayed in the output portion of the spreadsheet.

Detailed instructions for using the toolbox are discussed in the following sections. Table 5-4 at

the end of this chapter gives step by step instructions for using the toolbox.

5.2 Influent and Effluent

Default influent and effluent water characteristics data from Chapter 2 have been

embedded into the toolbox for users who do not know these parameters. The user may select and

use these data when running analyses by answering a series of questions at the top of the

spreadsheet. The series of questions and flow of logic through this part of the toolbox are

illustrated in Figure 5-1.

The user is first asked to input the water flow rate to the facility in bbl/day. The user is

then asked if the influent water chemistry is known. The user enters “yes” or “no” in the adjacent

cell (Use only lower case letters). The user then clicks the button that reads “Influent Data”. If

“yes” was entered the user knows the influent water chemistry, and the influent water constituent

values will be set to zero. The user then enters the known water chemistry (TDS is a required

input). The toolbox will approximate values for all constituent concentrations that are not known

except TDS. TDS must be entered for the toolbox to perform. To conduct the approximation, the

toolbox multiplies the entered TDS concentration by the ratio of the unknown constituent with

respect to the TDS based on the average concentrations of constituents in PRB CBM waters

(adapted from Table 2-1).
95

Figure 5-1: Influent and Effluent water characteristics series of questions and toolbox flow.

Do you know the

influent water

characteristics?

YES

NO

In what

watershed are the

wells located?

Input influent water

Characteristics

Do you want to use

average, high, or low

concentrations?

Influent water

Characteristics

Do you know the

desired effluent water

characteristics?

Input effluent water

Characteristics

YES

NO

•Surface Discharge (1)

•Drinking water (2)

•Irrigation (3)

•Livestock (4)

•EOR (5)

•Cooling Water (6)

Effluent Water

Characteristics

•Cheyenne (1)

•Belle Fourche (2)

•Little Powder (3)

•Powder (4)

•Tongue (5)

What is the water

flow rate, bbl/day?

How will the

water be used?

Do you want to

use more or less

stringent values?

•Less (1)

•More (2)

•Average (1)

•High (2)

•Low (3)
96

If the user does not know the influent water characteristics he/she enters “no” in the

adjacent cell. The user is then asked, “In what watershed are the wells located?” The user has the

option of selecting either the Cheyenne, Belle Fourche, Little Powder, Powder, or Tongue River

watershed by entering 1, 2, 3, 4, or 5 in the adjacent cell. The user is then asked “Do you want to

use average, high, or low concentrations?” These are the average, high, and low concentrations

for the different watersheds found in Table 2-1, and are selected by entering 1, 2, or 3 in the

adjacent cells. The user then clicks the influent data button which brings the appropriate

concentrations to the user interface.

Next, the user is asked, “Do you know the effluent water characteristics?” If the answer is

“yes”, the user clicks the effluent data button. This will zero all effluent concentrations so the

user can enter the effluent water requirements as non-zero and non-negative values. If a value is

blank, zero, or negative the toolbox will assume the constituent is not regulated and make the

value equal to the influent value. If the user answers “no” (effluent characteristics are unknown),

then the user is asked, “How will the water be used?” The user has the option of surface

discharge, drinking water, irrigation, livestock, enhanced oil recovery, or cooling water by

entering 1, 2, 3, 4, 5, or 6 into the adjacent cell. The user is then asked, “Do you want to use

more or less stringent values?” The user makes a selection by entering 2 or 1 in the adjacent cell,

and retrieves effluent values by hitting ENTER and clicking the effluent data button.

5.3 Technologies

5.3.1 Input

Each demineralization technology discussed in Chapter 3 has been given its own macro

that is constructed using equations developed in Chapter 3. Technologies include: reverse
97

osmosis (RO), electrodialysis reversal (EDR), high efficiency electro-pressure membranes

(HEEPM), ion exchange with hydrochloric acid regeneration (IX HCl), and ion exchange with

sulfuric acid regeneration (IX H2SO4). To analyze one of the included technologies hit ENTER

to ensure no cells are selected then click the button with the technology’s name on it.

The user has the option, but is not required to, input some parameters that may be

important to the costs associated with a technology. Optional user input includes: electrical cost

($/kW-hr), sulfuric acid cost ($/ton), hydrochloric acid cost ($/ton), facility life (yrs), ion

exchange run time (hrs), and brine disposal cost ($/bbl of brine). If the user does not input values

for these parameters the default values listed in Table 5-2 will be used.

Table 5-2: Toolbox default economic parameters.

Parameter Value Source

Electric $0.07/kW-hr Beagle (2007)

HCl

a

$225/ton Brainerd (2007)

H2SO4

a

$135/ton Brainerd (2007)

IX Run Time 20 hr Beardsley (1994)

Facility Life 20 yr

a HCl and H2SO4 are 36% and 98% pure by weight, respectively.

5.3.2 Output

The toolbox output is below the toolbox input on the spreadsheet. Toolbox output

includes: treated water, blended water and brine concentrations; feed, bypass, brine and final

flow rate; itemized operating costs, capital costs, and overall treatment cost in $/bbl of water

entering the facility.

Other output is displayed with input data. These output include a column that specifies if

the water must be treated to meet an effluent requirement for each constituent. Unknown

influent concentration approximations and unspecified effluent concentrations are output as well.
98

The toolbox calculates the treated water constituent concentrations based on the

technology’s treatment capabilities and the influent water concentrations. The final treated water

fraction () is determined for each constituent from the treated water concentrations and the

desired effluent concentrations. The limiting constituent (constituent that leads to the highest

final treated water fraction) or constituent for which the water is being blended to meet a desired

effluent concentration is then output. If the selected technology can not treat the water to the

desired effluent criteria “LIMIT!!!” will appear under the effluent exceeded column in the row of

the constituent for which the effluent requirement is not being achieved.

With the final treated water fraction known, flow rates are determined. The feed and

brine flow rates are determined with the technology’s water recovery rate (r) and other flow

rates. The flows output include: treated water flow rate, bypass water flow rate, feed flow rate,

brine flow rate, final flow rate, and initial treated fraction. Pretreatment residual flow is

neglected.

Treatment costs are then determined based on the feed flow rate. Amortized treatment

costs are output in both dollars per year and dollars per barrel of water entering the facility. Costs

that are output include: electrical, labor, chemical (which includes alkalinity removal),

membrane replacement, resin replacement, iron removal pretreatment, total operating, total

capital and total cost per barrel of water entering the facility. The capital cost per bbl of water

entering the facility is amortized assuming straight line depreciation over the lifetime of the

facility.
99

5.4 Brine

5.4.1 Input

Brine disposal / management costs are calculated separately from treatment costs. After

the user has selected a treatment technology, a brine management option must be selected

otherwise the brine management cost will not be included in the total treatment cost. A brine

management option is selected by clicking one of the buttons under “Click to Select Brine

Management”. The user has the option of selecting deep well injection, evaporation with

crystallization, or evaporation pond brine management.

There are optional user inputs for brine management. These are located at the top right of

the spreadsheet. Optional inputs include: deep well injection cost ($/bbl of brine), evaporation

rate (ft/yr), evaporation pond unit cost ($/ft

2

), brine facility life (yrs), and electrical cost ($/kWhr). Default values with references for optional input are outlined in Table 5-3.

Table 5-3: Toolbox default brine management economic parameters with source

Parameter Value Source

Deep Well Injection Cost $5.00/bbl of brine Miers (2007)

Evaporation Rate 4.06 ft/yr Urruty (2007)

Evaporation Pond Unit Cost $1.83/ft

2 Urruty (2007)

Brine Facility Life 5 yr

Electric $0.07/kW-hr Beagle (2007)

5.4.2 Output

Once a brine management option is selected (after a treatment technology has been

selected) the toolbox calculates the total cost to treat the water entering the facility. This is output

below the technology and brine output.

The toolbox also calculates different parameters for the brine management option

selected. The cost per bbl of brine ($/bbl) and cost per bbl of water entering the facility ($/bbl) is
100

the only output for deep well injection. Evaporation pond output includes: the required

evaporative area (ft

2

), the pond depth (assumed to be 1.5 ft (Mickley, 2001)), the evaporation

rate (bbl/yr), operating cost ($/bbl of brine), capital cost ($), total cost per bbl of brine ($/bbl) and

brine management cost per bbl of water entering the facility ($/bbl). Evaporation with

crystallization output includes: evaporator and crystallizer capital cost ($), and operating and

maintenance cost ($/bbl), in addition to, the total cost per bbl of brine ($/bbl) and brine cost per

bbl of water entering the facility ($/bbl). Step by step instructions for using the toolbox are

outlined in Table 5-4.

Table 5-4: Step by step instructions for using the toolbox.

1) Specify produced water flow rate (bbl/day)

2) Select influent water criteria

3) Click “Influent” button

4) Input influent characteristics if known (TDS is a required input)

5) Select effluent criteria

6) Click “Effluent” button

7) Input effluent characteristics if known

8) Input optional parameters if known

9) Click a technology button

10) Click a brine management button

11) Get output

• Total water treatment cost is output below the itemized technology and brine

management costs.

• Technical output is below the cost output.
101

6. SENSITIVITY ANALYSIS

6.1 Overview

Sensitivity analysis was used to determine which variable or variables have the largest

and smallest effect on the toolbox’s results. This type of analysis allows the user and developers

to better understand what some of the key parameters for each technology are when deciding

which technology to utilize. The sensitivity of a technology to a parameter that may change, such

as the unit price of electricity, may help the user determine with more certainty what technology

to choose to treat PRB CBM waters.

Design Expert software was used to conduct sensitivity analysis on the toolbox. This

software performs analysis through the use of factorial designs and analysis of variance

(ANOVA). These statistical tools allow for the determination of the variables which have the

greatest and smallest impact on the result output by the toolbox (Stat-Ease, 2007).

6.2 Sensitivity Analysis Setup

To perform sensitivity analysis a simulation experiment must first be developed to ensure

that all variables of interest are included. Once the variables are identified a high and low value

is assigned to each variable when utilizing a 2 level factorial design (Montgomery, 1997). This is

the type of factorial design utilized in this analysis. Once a high and low value is assigned for

each variable of interest, Design Expert develops the simulation experiment. The simulation

experiment is a table which lists all the variables and what their values should be set at for each

run of the toolbox. The toolbox’s response or result for each run is then recorded (total treatment
102

cost in $/bbl of Qi). The simulation experiment setups with results are illustrated in

Appendix D.

Table 6-1 illustrates the factors or variables that sensitivity analysis was performed with

and includes the high and low values that were used in analysis. Common variables that were

used for all technologies included: unit cost of electricity, facility life, brine disposal cost,

influent flow rate (Qi) and sodium removed. Technology specific factors included operating and

maintenance replacement costs for IX resin, RO membranes, EDR membranes, and HEEPM

membranes (low value is the assumed value (Chapter 3) and the high value is 2 times the

assumed value). Acid cost was the final factor included in analysis (high and low values are the

current cost of hydrochloric acid and sulfuric acid in the PRB, respectively).

Table 6-1: Factors used in sensitivity analysis with high and low values assigned.

Parameter Units RO IX HEEPM EDR

High Low High Low High Low High Low

Electricity $/kW-hr 0.21 0.07

a

0.21 0.07 0.21 0.07 0.21 0.07

Brine Disposal $/bbl Qbr 5

a

1 5 1 5 1 5 1

Facility Life Yr 10 5 10 5 10 5 10 5

Flow Rate bbl/day Qi 100,000 10,000 100,000 10,000 100,000 10,000 100,000 10,000

Sodium Removed mg/L 500 100 500 100 500 100 500 100

Membrane Cost $/bbl Qf 0.034 0.017

a

--- --- 0.0252 0.0126

a

0.00504 0.00252

a

Resin Cost $/bbl Qf --- --- 0.0106 0.0053

a

--- --- --- ---

Acid Cost

b

$/ton 250 135 250 135 250 135 --- ---

a Current cost in the PRB (Chapter 5) or assumed value (Chapter 3)

b Current cost of hydrochloric acid or sulfuric acid in the PRB (Chapter 5)

The sodium removed factor analysis was performed by holding the influent water

constituent concentrations constant for all analysis with Ci-Na and Ci-TDS equal to 600 mg/L and

1074 mg/L, respectively. The influent water chemistry is outlined in Appendix D Table D-5. The

effluent concentration for sodium was then set at 100 mg/L for 500 mg/L of sodium removed,

and 500 mg/L for 100 mg/L of sodium removed. Performing the analysis in this manner held all
103

influent concentrations constant, and forced sodium to be the limiting constituent because no

other effluent requirements were specified.

The factorial design for each technology was 2

7-1

for IX, RO, and HEEPM, and 2

6-1

for

EDR. These are fractional factorial designs. The exponent, 7 for example, identifies the number

of factors in the factorial design while the -1 signifies half as many runs to generate a result. A 2

7

factorial design is a full factorial design, and requires 128 runs to determine the effect of each

variable and interaction on the final result while a 2

7-1

requires 64 runs.

As a result of using fractional factorial designs, confounding occurs. Confounding is a

design technique for arranging a complete factorial experiment in blocks, where the block size is

smaller than the number of treatment combinations in one replicate (Montgomery, 1997).

Confounding typically causes the effects of higher order interactions to be indistinguishable

(Montgomery, 1997).

The end result of confounding in a 2

6-1

fractional factorial design is that some three and

all four factor interactions can not be defined because they are aliased with other interactions

making them indistinguishable from other interactions. Likewise, a 2

7-1

fractional factorial

design can not define the effects of greater than a four factor interaction. However, the effects of

confounding are insignificant for this sensitivity analysis because first and second order

interactions are of the most concern. Therefore, fractional factorial designs have been utilized.

6.3 Results and Discussion

6.3.1 Background

After performing the simulation experiments by running the toolbox 64 times for all 2

7-1

factorial designs and 32 times for 2

6-1

factorial designs the responses (total treatment cost, $/bbl)
104

were entered into Design Expert. Design Expert then performed factorial design and ANOVA

analysis to aid in determining what factors are most important for each technology. The first

results output were the sum of squares and the relative percent contribution of each factor to the

total sum of squares. The factor with the largest percent contribution to the total sum of squares

has the largest effect on the toolbox’s response and vice versa (Montgomery, 1997).

The relative contribution of each factor to the toolbox’s response can also be seen with

the use of a perturbation plot (Figure 6-1A) or normal plot (Figure 6-1B), both of which are

developed by Design Expert. The perturbation plot (Figure 6-1A) is a graph with response on the

y-axis and the factor level on the x-axis (-1 and 1 correspond to the low and high values,

respectively). This plot helps to compare the effect of all the factors at a particular point in the

design space. The response is plotted by changing only one factor over its range, while holding

the other factors constant. The factor whose line has the largest slope or curvature has the largest

effect and vise versa on the perturbation plot (Stat-Ease, 2007).

A) Perturbation

Factor Level

Treatment Cost, $/bbl

-1 0 1

0.01

0.06

0.11

0.16

0.21

A

B A

B

C

C D

D

E

E

F

F

G

G

A) Perturbation

Factor Level

Treatment Cost, $/bbl

-1 0 1

0.01

0.06

0.11

0.16

0.21

A

B A

B

C

C D

D

E

E

F

F

G

G

A: Acid Cost

B: Membrane Cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: flow

G: Electric

Positive Effects

Negative Effects

Normal %

Probability

Standardized Effect

B) Normal Plot

-0.01 0.02 0.05 0.08 0.11

1

5

10

20

30

50

70

80

90

95

99

A

B

C

D

E

F

G

AE

CE

DE

EF

EG A: Acid Cost

B: Membrane Cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: flow

G: Electric

Positive Effects

Negative Effects

Normal %

Probability

Standardized Effect

B) Normal Plot

-0.01 0.02 0.05 0.08 0.11

1

5

10

20

30

50

70

80

90

95

99

A

B

C

D

E

F

G

AE

CE

DE

EF

EG

Figure 6-1: Reverse osmosis sensitivity results A) Perturbation Plot and B) Normal Plot.
105

A normal plot identifies each factor and multiple factor interactions as data points on a

graph with normal percent probability on the y-axis and standardized effect on the x-axis. The

factor with the greatest contribution is the farthest from a standard effect equal to zero and

factors with the least contribution, residuals, are closest to zero standard effect.

All main factors and multiple factor interactions with greater than or equal to 1%

contribution to the total sum of squares were selected on a normal plot output in Design Expert.

Selected factors were then included in the ANOVA analysis. Interactions that were not included

are residuals.

6.3.2 Reverse Osmosis

Sensitivity results for reverse osmosis are presented in Table 6-2 and Figure 6-1. RO

treatment cost is most sensitive to the amount of sodium removed which accounts for 84.1% of

the total sum of squares. This is also illustrated in Figures 6-1 A and B as the line representing

sodium removed has the greatest slope (Figure 6-1 A) and sodium removed has the largest

standardized effect (Figure 6-1 B).

Table 6-2: Reverse osmosis factorial design and ANOVA results. df is the degrees of freedom.

Total degrees of freedom equals 63.

Term Effect Sum of Squares % Contribution F-value Prob > F

A-Acid Cost 0.00734 0.000863 0.37 60 < 0.0001

B-Membrane Cost -0.00528 0.000446 0.19 31 < 0.0001

C-Facility Life -0.01191 0.002268 0.98 157 < 0.0001

D-Brine Disposal 0.03297 0.017391 7.54 1,203 < 0.0001

E-Sodium Removed 0.11016 0.194150 84.1 13,425 < 0.0001

F-Flow -0.01028 0.001691 0.73 117 < 0.0001

G-Electric 0.01409 0.003178 1.38 220 < 0.0001

DE 0.02159 0.007461 3.23 516 < 0.0001

EG 0.00959 0.001473 0.64 102 < 0.0001

CE -0.00553 0.000490 0.21 34 < 0.0001

AE 0.00459 0.000338 0.15 23 < 0.0001

EF -0.00378 0.000229 0.10 16 0.0002

Residuals (df = 51) 0.00053 0.000738 0.32

Total 0.16409 0.230715 100
106

Brine disposal cost is the second most sensitive variable and accounts for 7.54% of the

total sum of squares. There is also an interaction effect of sodium removed and brine disposal

cost. This interaction is attributed to more sodium removal generating more brine and accounts

for 3.23% of the total sum of squares. Other main factors, interaction effects, and residuals have

a minimal contribution accounting for 5.07% of the total sum of squares as a whole. However,

all terms presented in Table 6-2 are statistically significant because the Prob > F is less than 0.05

(Stat-Ease, 2007).

6.3.3 Ion Exchange

Sensitivity results for ion exchange are presented in Table 6-3 and Figure 6-2. IX results

show sodium removed having the largest contribution to response which accounts for

81.9% of the total sum of squares. This is also illustrated in Figures 6-2 A and B.

Table 6-3: Ion exchange factorial design and ANOVA results. df is the degrees of freedom.

Total degrees of freedom equals 63.

Term Effect Sum of Squares % Contribution F-value Prob > F

A-Acid Cost 0.03216 0.016544 6.20 46,661 < 0.0001

B-Resin cost -0.00128 0.000026 0.01 74 < 0.0001

C-Facility Life -0.01141 0.002082 0.78 5,871 < 0.0001

D-Brine Disposal 0.02872 0.013196 4.94 37,218 < 0.0001

E-Sodium Removed 0.11691 0.218673 81.9 616,732 < 0.0001

F-flow -0.01303 0.002717 1.02 7,663 < 0.0001

G-Electric 0.00003 0.000000 0.00 0 0.8345

AE 0.02141 0.007332 2.75 20,678 < 0.0001

DE 0.01909 0.005833 2.18 16,451 < 0.0001

CF 0.00441 0.000311 0.12 876 < 0.0001

CE -0.00403 0.000260 0.10 733 < 0.0001

Residuals (df = 52) -0.00113 0.000018 0.01

Total 0.19184 0.266993 100

Acid cost is the second most significant factor for IX and accounts for 6.2 % of the total

sum of squares. There is also an interaction effect (Term AE) associated with both sodium
107

removed and acid cost because more sodium removed requires more acid use. This term

accounts for 2.75% of the total sum of squares.

The next most significant factor is brine disposal cost which accounts for 4.94% of the

total sum of squares. This factor also has an interaction effect with sodium removed (Term DE)

accounting for 2.18% of the total sum of squares. These factors cause an interaction effect

because more sodium removed produces more brine as a result of more regenerant or acid being

required.

All other factors combined account for 2.08% of the total sum of squares. All terms

presented in Table 6-3 are statistically significant with the exception of electrical price which has

a Prob > F equal to 0.8345.

Factor Level

Treatment Cost, $/bbl

A) Perturbation

-1 0 1

0.01

0.065

0.12

0.175

0.23

A

A

B B

C

C D

D

E

E

F

F

G

G

Factor Level

Treatment Cost, $/bbl

A) Perturbation

-1 0 1

0.01

0.065

0.12

0.175

0.23

A

A

B B

C

C D

D

E

E

F

F

G

G

A: Acid Cost

B: Resin cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: Flow

G: Electric

Positive Effects

Negative Effects

Normal %

Probability

Standardized Effect

B) Normal Plot

-0.01 0.02 0.05 0.08 0.12

1

5

10

20

30

50

70

80

90

95

99

A

B

C

D

E

F

G

AE

CE

CF DE

A: Acid Cost

B: Resin cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: Flow

G: Electric

Positive Effects

Negative Effects

Normal %

Probability

Standardized Effect

B) Normal Plot

-0.01 0.02 0.05 0.08 0.12

1

5

10

20

30

50

70

80

90

95

99

A

B

C

D

E

F

G

AE

CE

CF DE

Figure 6-2: Ion exchange sensitivity results A) Perturbation Plot and B) Normal Plot.
108

6.3.4 Electrodialysis Reversal

Sensitivity results for electrodialysis reversal are presented in Table 6-4 and Figure 6-3.

The most sensitive factor for EDR treatment cost is facility life which accounts for 55.6% of the

total sum of squares. This is in contrast to results presented above for RO and IX, and is due to

capital costs accounting for most of the costs associated with EDR. The effect of the variable is

also negative which means that if EDR is utilized the facility life should be maximized to

minimize overall treatment cost.

The second most sensitive variable for EDR is influent flow rate (Qi) which accounts for

29.8% of the total sum of squares. This factor also has a negative effect which means that the

influent flow rate for an EDR facility should be maximized to minimize treatment costs.

Table 6-4: Electrodialysis reversal factorial design and ANOVA results. df is the degrees of

freedom. Total degrees of freedom equals 31.

Term Effect Sum of Squares % Contribution F-value Prob > F

A-Electric 0.01206 0.001164 3.12 10,506 < 0.0001

B-Membrane Cost -0.00406 0.000132 0.35 1,192 < 0.0001

C-Facility Life -0.05094 0.020757 55.6 187,346 < 0.0001

D-Brine Disposal 0.01469 0.001726 4.62 15,576 < 0.0001

E-Sodium Removed 0.01119 0.001001 2.68 9,037 < 0.0001

F-Flow -0.03731 0.011138 29.8 100,526 < 0.0001

CF 0.01244 0.001238 3.31 11,170 < 0.0001

DE 0.00406 0.000132 0.35 1,192 < 0.0001

AE 0.00244 0.000048 0.12 429 < 0.0001

Residuals (df = 22) -0.00013 0.000002 0.01

Total -0.03556 0.037337 100

Both facility life and influent flow rate have an interaction effect (Term CF) that accounts

for 3.31% of the total sum of squares. This effect shows that if EDR is used to treat PRB CBM

waters, the facility life and influent flow rate should be maximized.
109

A) Perturbation

Factor Level

Treatment Cost, $/b

bl

-1 0 1

0.07

0.1025

0.135

0.1675

0.2

A

A

B

B

C

C

D

D

E

E

F

F

A) Perturbation

Factor Level

Treatment Cost, $/b

bl

-1 0 1

0.07

0.1025

0.135

0.1675

0.2

A

A

B

B

C

C

D

D

E

E

F

F

A: Electric

B: Membrane Cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: Flow

Positive Effects

Negative Effects

Normal %

Probability

Standardized Effect

B) Normal Plot

-0.05 -0.03 -0.02 0 0.01

1

5

10

20

30

50

70

80

90

95

99

A

B

C

D

E

F

AE

CF

DE

A: Electric

B: Membrane Cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: Flow

Positive Effects

Negative Effects

Normal %

Probability

Standardized Effect

B) Normal Plot

-0.05 -0.03 -0.02 0 0.01

1

5

10

20

30

50

70

80

90

95

99

A

B

C

D

E

F

AE

CF

DE

Figure 6-3: Electrodialysis reversal sensitivity results A) Perturbation Plot and B) Normal Plot.

Other factors that the toolbox’s response is sensitive to include brine disposal cost,

electrical cost and sodium removed, each accounting for 4.62%, 3.12% and 2.68% of the total

sum of squares, respectively. The remaining factor (membrane replacement cost) and interaction

effects have minimal effect accounting for 0.83% of the total sum of squares as a whole. All

factors and interaction effects are statistically significant though.

6.3.5 HEEPM

Sensitivity results for HEEPM are presented in Table 6-5 and Figures 6-4 A and B.

HEEPM treatment cost is most sensitive to sodium removed which accounts for 80.7% of the

total sum of squares.
110

Table 6-5: HEEPM factorial design and ANOVA results. df is the degrees of freedom. Total

degrees of freedom equals 63.

Term Effect Sum of Squares % Contribution F-value Prob > F

A-Acid Cost 0.00853 0.001165 0.24 6 0.0187

B-Membrane Cost -0.00603 0.000582 0.12 3 0.0919

C-Facility Life -0.02441 0.009531 1.94 48 < 0.0001

D-Brine Disposal 0.01522 0.003706 0.76 19 < 0.0001

E-Sodium Removed 0.15734 0.396113 80.7 2,008 < 0.0001

F-Flow -0.01953 0.006104 1.24 31 < 0.0001

G-Electric 0.04391 0.030844 6.29 156 < 0.0001

EG 0.03528 0.019916 4.06 101 < 0.0001

CE -0.01778 0.005059 1.03 26 < 0.0001

EF -0.01403 0.003150 0.64 16 0.0002

CF 0.01209 0.002340 0.48 12 0.0012

AE 0.01116 0.001991 0.41 10 0.0025

Residuals (df = 51) 0.00478 0.010058 2.05

Total 0.20653 0.490558 100

The next most significant factor for HEEPM is the cost of electricity which accounts for

6.29% of the total sum of squares. There is also an interaction effect of electrical cost and

sodium removed (Term EG) that accounts for 4.06% of the total sum of squares.

Facility life and influent flow rate are the next most significant factors and account for

1.94% and 1.24% of the total sum of squares, respectively. These factors have a negative effect

meaning that HEEPM may be more cost effective with a long facility life and higher influent

flow rate. However, the interaction effect of these two terms (Term CF) accounts for only 0.48%

of the total sum of squares.

Other factors and interaction effects have little effect on toolbox response and account for

5.73% of the total sum of squares as a whole. However, all factors and interaction effects

included in the analysis are statistically significant with Prob < F less than 0.05.
111

A: Acid Cost

B: Membrane Cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: Flow

G: Electric

Positive Effects

Negative Effects

B) Normal Plot

Normal %

Probability

Standardized Effect

-0.02 0.02 0.07 0.11 0.16

0.1

1.0

5.0

10.0

20.0

30.0

50.0

70.0

80.0

90.0

95.0

99.0

99.9

A

B

C

D

E

F

G

AE

CE

CF

EF

EG

A: Acid Cost

B: Membrane Cost

C: Facility Life

D: Brine Disposal

E: Sodium Removed

F: Flow

G: Electric

Positive Effects

Negative Effects

B) Normal Plot

Normal %

Probability

Standardized Effect

-0.02 0.02 0.07 0.11 0.16

0.1

1.0

5.0

10.0

20.0

30.0

50.0

70.0

80.0

90.0

95.0

99.0

99.9

A

B

C

D

E

F

G

AE

CE

CF

EF

EG

A) Perturbation

Factor Level

Treatment Cost, $/b

bl

-1 0 1

0.02

0.095

0.17

0.245

0.32

A

B A

B

C

C

D

D

E

E

F

G

G

F

A) Perturbation

Factor Level

Treatment Cost, $/b

bl

-1 0 1

0.02

0.095

0.17

0.245

0.32

A

B A

B

C

C

D

D

E

E

F

G

G

F

Figure 6-4: HEEPM sensitivity results A) Perturbation Plot and B) Normal Plot.

6.4 Summary

The sensitivity of total treatment cost for all of the selected technologies to key variables

have been determined through the use of Design Expert software. Results have shown that the

total treatment cost is most sensitive to sodium removed when utilizing RO, IX, and HEEPM.

EDR treatment cost is most sensitive to facility life and influent flow rate.

Second order interactions and other variables such as brine disposal cost, acid cost,

facility life and influent flow have been shown to have an effect on the treatment cost for all

technologies, but to a lesser degree than the above listed variables. Technology specific

parameters such as membrane and resin replacement costs have been shown to have the least

effect on treatment costs.
112

7. TECHNOLOGY COST EFFECTIVENSS AND SELECTION

7.1 Overview

The purpose of this chapter is to show toolbox results for all of the selected technologies

in a comparative manner. This analysis may aid an individual in deciding what technology to

select and under what conditions when treating CBM waters.

Conditions that have been tested over a range of sodium removals include: the anticipated

treatment facility life, influent flow rate, brine disposal cost, electrical cost, and process chemical

cost. These variables were selected because they were shown in Chapter 6 to have the greatest

impact on the total treatment cost for selected technologies.

7.2 Scenarios

The following analysis was conducted to determine the impact of key variables on the

total treatment cost ($/bbl) when deciding what technology to use to treat PRB CBM waters. To

do this, optional user inputs (key variables) were changed one at a time for a total of seven

scenarios (Table 7-1) and the toolbox was run with six different sodium removals (Table 7-2) per

scenario to determine what technology may be more cost effective over another under certain

conditions.

Sodium removed was selected as the performance parameter because the total treatment

cost of each technology, excluding EDR, is most sensitive to sodium removed. Treatment cost

($/bbl) versus sodium removed was then plotted for selected treatment technologies to determine

under what conditions one technology may be more cost effective than another.
113

Table 7-1: Scenarios Used to Evaluate Technology Costs and Cost Effectiveness. “x” and “y”

equal the same value used in scenario 1 and 2, respectively.

Scenario

Parameter Units

1 2 3 4 5 6 7

Treatment Facility Life yrs 5 10 y y y y y

Influent Flow Rate bbl/day 100,000 x 10,000 x x x x

Brine Disposal $/bbl brine 5 x x 1 x x x

Electric Cost $/kW-hr 0.07 x x x 0.21 x x

HCl Cost $/ton 225 x x x x 135 x

H2SO4 Cost $/ton 135 x x x x x 225

For these analyses the effluent sodium concentration was held constant at 100mg/L and

the influent concentration was changed by inputting a different influent TDS concentration for

each of the six simulations for each scenario (Table 7-2). The toolbox then calculated the influent

sodium concentration based on the influent TDS concentration. Influent TDS (user defined and

toolbox calculated), toolbox-predicted influent sodium, and sodium removed for these analyses

are outlined in Table 7-2.

Table 7-2: Water Parameters Used to Evaluate Technology Costs and Cost Effectiveness. All

units are mg/L.

User FTDS

a Calculated FTDS

b

Feed Na

b Effluent Naa Na Removedc

2500 2774 858 100 758

2000 2219 687 100 587

1500 1665 515 100 415

1000 1110 343 100 243

750 832 257 100 157

500 555 172 100 72

a

Selected Input

b Determined by toolbox

c

Feed Na – Effluent Na

Other specified inputs were facility life, influent flow rate, brine disposal cost, cost of

electricity, hydrochloric acid cost and sulfuric acid cost. Values used for the seven different

scenarios tested are outlined in Table 7-1. Facility life and flow rates were arbitrarily selected
114

and may be representative of current and/or future operations in the PRB. Brine disposal costs

that were selected represent the current cost of deep well injection brine disposal in the PRB

($5/bbl) and what may happen if this cost is reduced ($1/bbl). The costs selected for electricity

represent the current cost in the PRB ($0.07/kW-hr) and the tripling of the current cost of

electricity ($0.21/kW-hr). Acid costs represent the current hydrochloric acid ($225/ton) and

sulfuric acid cost ($135/ton) in the PRB.

7.3 Results and Discussion

7.3.1 Scenarios 1 and 2: Facility Life

The facility life has been shown to be an important parameter with respect to the total

treatment cost when utilizing EDR. This parameter is important because the toolbox assumes

straight line depreciation of the treatment technology’s capital cost over the life of the facility.

Therefore, an increased facility life will reduce the treatment cost per bbl of influent water. This

relationship has been tested to determine which technology may be the most cost effective for a

facility life equal to 5yrs (Figure 7-1) and 10yrs (Figure 7-3). Default values were used for all

analyses unless otherwise noted (see Table 7-1). The intersection of two or more lines on these

plots is considered a crossover point.

When the treatment facility life is 5 yrs IX H2SO4 is the most cost effective treatment

technology for lower amounts of sodium removed (sodium removed<760 mg/L). However, when

the sodium removed is above 760 mg/L, EDR becomes more cost effective (at cross over point

shown in Figure 7-1).

There are safety issues related to the use of 98% pure sulfuric acid in the PRB. The

primary issue is the corrosiveness of 98% pure sulfuric acid. This makes shipping, storage and
115

handling more difficult than what is required for 36% pure hydrochloric acid. Another issue is

the possibility of precipitate formation (BaSO4 or CaSO4) which may cause adverse effects to

process pumps or vessels. For these reasons, it may be necessary to choose IX HCl which is the

next most cost effective technology, for sodium removed< 360 mg/L. When sodium removed is

greater than 360 mg/L, EDR becomes more cost effective than IX HCl.

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-1: Scenario 1, facility life equal to 5 years, treatment cost vs. sodium removed (Refer

to Table 7-1 for full Scenario description). Arrows represent crossover points discussed in the

text.

Figure 7-2 illustrates itemized costs associated with selected technologies, and Table 7-3

outlines the percent contribution of each cost to the overall treatment cost when removing 243

mg/L of sodium. Note that the highest contributor to treatment cost for IX HCl and IX H2SO4 is

chemical cost (53%) and brine disposal cost (48%), respectively, but the brine disposal cost

($0.034/bbl) for each is equal because each generates equal volumes of brine. This difference

exists because hydrochloric acid costs $0.025/regenerative equivalent while H2SO4 costs
116

$0.007/regenerative equivalent for the conditions of Scenario 1. Therefore, decreased safety

concerns come at the price of a more expensive acid (HCl).

While IX is the most cost effective technology for lower amounts of sodium removed,

EDR is the most cost effective technology when higher amounts of sodium must be removed.

The number one contributor to the cost of utilizing EDR is capital cost. This accounts for 64% of

the total treatment cost of $0.149/bbl of influent water (Table 7-3).

$0.00

$0.02

$0.04

$0.06

$0.08

$0.10

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otalTreatment C

ost, $/b

bl

Figure 7-2: Scenario 1 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for Scenario description)

Table 7-3: Scenario 1: Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed. (Refer to Table 7-1 for Scenario description)

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 12% 19% 64% 34% 17%

Brine 30% 48% 14% 14% 42%

Electric 0.1% 0.1% 4% 23% 6%

Labor 1.2% 1.9% 12% 7.4% 6%

Chemical 53% 25% 0.4% 15% 15%

Parts 3% 5% 3% 6% 11%

Iron Pretreatment 0% 0% 3% 2% 2%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.111 0.069 0.149 0.148 0.147
117

Figures 7-3 and 7-4, and Table 7-4 show toolbox results if the facility life is doubled and

equal to ten years. Results show that IX with sulfuric acid regeneration is still the most cost

effective for lower amounts of sodium removed (sodium removed<540 mg/L), but the crossover

point where EDR becomes more cost effective has changed to sodium removed greater than 540

mg/L. This change is a direct result of capital cost driving the cost of EDR treatment. Note that

capital cost is still the main contributor to EDR treatment cost, but to a lesser extent (47%).

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-3: Scenario 2, facility life equal to 10 years, treatment cost vs. sodium removed (Refer

to Table 7-1 for full Scenario description). Arrows represent crossover points discussed in the

text.

The impact of a longer treatment facility life has also impacted the capital cost percent

contribution of all technologies. The percent contribution of capital cost to the overall treatment

cost has decreased for all technologies when comparing Table 7-3 (life = 5 yrs) to Table 7-4 (life

= 10 yrs). The result is a 6%, 8%, 17%, 14%, and 8% decrease (versus total) in the contribution

of capital cost for IX HCl, IX H2SO4, EDR, HEEPM, and RO, respectively.
118

$0.00

$0.02

$0.04

$0.06

$0.08

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otalTreatment C

ost, $/b

bl

Figure 7-4: Scenario 2 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for full Scenario description)

Table 7-4: Scenario 2 Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed. (Refer to Table 7-1 for full Scenario description)

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 6% 11% 47% 20% 9%

Brine 32% 54% 20% 16% 46%

Electric 0.1% 0.1% 6.4% 27.0% 7.0%

Labor 1% 2% 17% 9% 7%

Chemical 56% 27% 1% 18% 16%

Parts 4% 6% 4% 7% 12%

Iron Pretreatment 0% 0% 5% 3% 3%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.104 0.063 0.102 0.124 0.135

These reductions in percent contribution of capital cost have resulted in a decrease in

overall treatment cost for all technologies (Tables 7-3 and 7-4). Therefore, the facility life should

be maximized regardless of the technology selected to treat PRB CBM waters. Note that costs
119

presented do not include interest rates or salvage values; both will affect the actual treatment

cost.

7.3.2 Scenario 3: Influent Flow Rate

Influent flow rate significantly affects EDR costs and is also important for other

technologies. Economies of scale become significant at higher flow rates. Therefore, an

increased influent flow rate will reduce the treatment cost per bbl of influent water. This

relationship has been tested to determine which technology may be the most cost effective for a

10 year facility life with an influent flow rate equal to 100,000 bbl/day (Figure 7-3) versus

10,000 bbl/day (Figure 7-5).

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-5: Scenario 3, influent flow rate equal to 10,000bbl/day and 10 year facility life,

treatment cost vs. sodium removed (Refer to Table 7-1 for full Scenario description). Arrows

represent crossover points discussed in the text.

The crossover point between IX H2SO4 and EDR has increased from sodium removed

equal to approximately 540 mg/L (Figure 7-3) to 715 mg/L (Figure 7-5). This increase is a direct

result of facility size, and can be seen by comparing Figures 7-4 and 7-6 and Tables 7-4 and 7-5.
120

Capital costs per barrel of influent water and the percent contribution of capital cost to total

treatment cost has increased as a result of a lower influent flow rate.

$0.00

$0.02

$0.04

$0.06

$0.08

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otalTreatment C

ost, $/b

bl

Figure 7-6: Scenario 3 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for full Scenario description)

Table 7-5: Scenario 3: Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed. (Refer to Table 7-1 for full Scenario description)

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 14% 22% 59% 32% 15%

Brine 30% 47% 16% 14% 43%

Electric 0.1% 0.1% 5% 23.1% 6.6%

Labor 1% 2% 13% 8% 6%

Chemical 52% 24% 0.4% 15% 15%

Parts 3% 5% 3% 6% 12%

Iron Pretreatment 0% 0% 4% 2% 2%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.113 0.071 0.132 0.145 0.143

The lower flow rate does not allow for economies of scale, but the higher flow rate does.

This is evident because the percent contribution of capital cost (versus total) has increased by
121

8%, 11%, 12%, 12%, and 6% for IX HCl, IX H2SO4, EDR, HEEPM, and RO, respectively

(Table 7-4 vs. Table 7-5). Therefore, to minimize treatment costs, regardless of the technology

selected, the influent flow rate should be maximized. However, costs such as pipelines have not

been included, so the total cost may be higher, for the larger facility, if the facility is not placed

in a centralized location with respect to the CBM wells from which water is being treated. Also,

note that a decrease in facility size does yield a lower capital investment.

7.3.3 Scenario 4: Brine Disposal Cost

Brine disposal cost has been shown to be significant with respect to the total treatment

cost for all technologies. To determine how brine management/disposal cost effects crossover

points two brine disposal costs were examined ($1/bbl and $5/bbl of brine).

Figure 7-3 illustrates the treatment cost associated with selected technologies when the

facility life is ten years and brine disposal cost is $5/bbl of brine (the default value). Figure 7-7

illustrates treatment costs when the brine disposal cost is reduced to $1 /bbl of brine. The influent

flow rate and facility life were held constant and equal to 100,000 bbl/day and 10 yrs,

respectively.

When comparing the results of Figure 7-3 to Figure 7-7 it is evident that crossover points

change. Most notably, the cross over point where EDR becomes the most cost effective has

increased from approximately 540 mg/L of sodium removed to over 760 mg/L. This increase is

because EDR generates much less brine than the other technologies making EDR less cost

effective at lower brine disposal costs.

Also, RO is no longer the most costly treatment technology. This is because RO

generates the most brine of all the technologies which makes RO treatment costs very sensitive

to brine disposal cost. For example, if the brine disposal cost were reduced to $0.50/bbl of brine
122

for 243 mg/L of sodium removed the total treatment costs would be $0.074, $0.032, $0.083,

$0.105 and $0.079/bbl for IX HCl, IX H2SO4, EDR, HEEPM and RO, respectively. This change

makes RO nearly the same cost as IX HCl. This means that RO may be selected if IX H2SO4 is

not selected due to safety concerns. Brine disposal costs clearly influence what technology is the

most cost effective.

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-7: Scenario 4, brine disposal cost equal to $1/bbl of brine, treatment cost vs. sodium

removed (Refer to Table 7-1 for full Scenario description). Arrows represent crossover points

discussed in the text.

Figure 7-8 and Table 7-6 illustrate the itemized costs and percent contribution of each

cost, respectively, when removing 243 mg/L of sodium from PRB CBM waters when the brine

disposal cost is $1/bbl of brine. This change in brine disposal cost from $5/bbl has resulted in a

decrease in overall treatment cost for all of the technologies (Table 7-4 versus Table 7-6) and an

overall decrease in the percent contribution of brine disposal costs to total treatment cost (Table

7-4 vs. Table 7-6 and Figure 7-4 vs. Figure 7-8).
123

$0.00

$0.02

$0.04

$0.06

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otalTreatment C

ost, $/b

bl

Figure 7-8: Scenario 4 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for full Scenario description)

Table 7-6: Scenario 4 Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed. (Refer to Table 7-1 for full Scenario description)

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 9% 19% 56% 23% 15%

Brine 9% 19% 5% 4% 14%

Electric 0.1% 0.3% 7.7% 31% 11%

Labor 2% 4% 21% 10% 10%

Chemical 76% 48% 1% 20% 26%

Parts 5% 11% 5% 8% 19%

Iron Pretreatment 0% 0% 5% 3% 4%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.078 0.036 0.085 0.107 0.085

The sensitivity of EDR treatment cost to EDR labor cost was investigated after running

these analyses because labor cost was found to contribute 21% to EDR total treatment cost

(Table 7-6). Results of this sensitivity analysis (Table D-7) showed labor to contribute 1.3% to
124

the total sum of squares. However, the main contributors to the total sum of squares did not

change significantly (Table D-7 vs. Table 6-4).

7.3.4 Scenario 5: Electrical Cost

The total treatment costs for all selected technologies, except IX, showed sensitivity to

the cost of electricity. Analysis was therefore conducted to determine how the cost of electricity

affects crossover points and cost effectiveness. This was done with two different electrical costs

$0.07/kW-hr (the default value) and $0.21/kW-hr for an influent flow rate equal to 100,000

bbl/day and treatment facility life equal to 10 yrs. Results are illustrated in Figures 7-3 and 7-9.

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-9: Scenario 5, electrical cost equal to $0.21/kW-hr, treatment cost vs. sodium removed

(Refer to Table 7-1 for full Scenario description). Arrows represent crossover points discussed in

the text.

Increasing the cost of electricity increases all cross over points. Most notably, EDR does

not become the most cost effective until the sodium removed is above 280 mg/L (IX HCl) and

over 760 mg/L (IX H2SO4). Also, the RO, HEEPM and IX HCl lines have become nearly

parallel as a result of increasing the cost of electricity.
125

Figure 7-10 and Table 7-7 illustrate itemized costs and percent contribution of each cost

component to the overall treatment cost, respectively, for sodium removed equal to 243 mg/L.

This increase in electrical cost has caused the percent contribution of electricity to increase for

all technologies (Table 7-4 vs. Table 7-7).

$0.00

$0.02

$0.04

$0.06

$0.08

$0.10

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otalTreatment C

ost, $/b

bl

Figure 7-10: Scenario 5 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for full Scenario description)

Also the electrical cost component for all technologies has increased. Most notably, the

electrical component for HEEPM and RO increased from $0.035 (Figure 7-4) to $0.10/bbl

(Figure 7-10) and $0.01 to $0.03/bbl, respectively. However, electrical cost does not appear to

have a large influence on overall technology cost effectiveness because IX is not an energy

intensive process, but if energy conserving processes or producing processes (energy recovery

turbines or solar power) were utilized the electrical cost component would decrease which may

change cost effectiveness.
126

Table 7-7: Scenario 5 Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed. (Refer to Table 7-1 for full Scenario description)

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 6% 11% 41% 13% 8%

Brine 32% 54% 18% 11% 40%

Electric 0.3% 0.4% 17% 52.6% 18.4%

Labor 1% 2% 15% 6% 6%

Chemical 56% 27% 1% 11% 14%

Parts 4% 6% 4% 5% 11%

Iron Pretreatment 0% 0% 4% 2% 2%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.105 0.063 0.106 0.190 0.153

7.3.5 Scenarios 6 and 7: Acid Cost

The final factor that was tested is acid cost. Two acid costs were selected for this analysis

$135/ton and $225/ton (default values for hydrochloric acid and sulfuric acid). The cost for each

type of acid was set equal to each other on a per ton basis, the facility life and influent flow rates

were then set equal to 10 yrs and 100,000 bbl/day, respectively. Results are illustrated in Figures

7-11 and 7-12.

The first thing to note from Figures 7-11 and 7-12 is that the IX HCl and IX H2SO4 lines

are not the same. This is a direct result of the concentration of each type of acid. The bulk

concentration of sulfuric acid and hydrochloric are assumed to be 98% and 36% acid by weight.

So, even though the acid cost is the same for each type of acid the treatment cost is still different.

When comparing results from Figures 7-11 and 7-12 crossover points change as a result

of acid cost. The crossover points at the lower acid cost of $135/ton (Figure 7-11) shows EDR

being the most cost effective when greater than 350 mg/L (IX HCl) and 540 mg/L (IX H2SO4) of

sodium are removed. The higher acid cost of $225/ton illustrated in Figure 7-12 makes EDR the

most cost effective treatment technology when greater than 235 (IX HCl) and 410 mg/L (IX
127

H2SO4) of sodium are removed. Therefore the cost of acid does influence technology cost

effectiveness.

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-11: Scenario 6, acid cost equal to $135/ton, treatment cost vs. sodium removed (Refer

to Table 7-1 for full Scenario description). Arrows represent crossover points discussed in the

text.

Also, note that RO is the most cost prohibitive when greater than 160 mg/L of sodium is

removed for these analyses. This is most likely attributed to the default brine disposal cost of

$5/bbl of brine being used for these analyses.

Figures 7-13 and 7-14, and Tables 7-8 and 7-9 illustrate itemized treatment costs and the

percent contribution of each cost to the total treatment cost, respectively, for an acid cost of $135

and $225/ton. As a result of increasing acid cost, the percent contribution of chemical cost

increases for all technologies except EDR (Table 7-8 vs. Table 7-9). EDR does not change

because chemical/acid requirements are considered constant for this process.
128

0

0.1

0.2

0.3

0.4

0 100 200 300 400 500 600 700 800

T

r

e

atm

e

nt C

ost, $/b

bl

Sodium Removed, mg/L

IX HCl IX H2SO4 EDR HEEPM RO

Figure 7-12: Scenario 7, acid cost equal to $225/ton, treatment cost vs. sodium removed (Refer

to Table 7-1 for full Scenario description). Arrows represent crossover points discussed in the

text.

$0.00

$0.02

$0.04

$0.06

$0.08

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otalTreatment C

ost, $/b

bl

Figure 7-13: Scenario 6 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for full Scenario description).
129

Treatment cost changes are the most pronounced for both IX processes, RO and HEEPM.

This is because the chemical/acid requirements are dependent on treatment requirements for

these technologies. However, acid cost does not appear to have a significant affect on cost

effectiveness when the sodium removed is less than 410 mg/L (Figure 7-12).

Table 7-8: Scenario 6 Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed. (Refer to Table 7-1 for full Scenario description)

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 8% 11% 47% 20% 9%

Brine 42% 54% 20% 16% 46%

Electric 0.1% 0.1% 6% 27% 7%

Labor 2% 2% 17% 9% 7%

Chemical 44% 27% 1% 18% 16%

Parts 5% 6% 4% 7% 12%

Iron Pretreatment 0% 0% 5% 3% 3%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.081 0.063 0.102 0.124 0.135

$0.00

$0.02

$0.04

$0.06

$0.08

IX HCl IX H2SO4 EDR HEEPM RO

Capital Brine Electric

Labor Chemical Parts

Iron Pretreatment

T

otal Treatment C

ost, $/b

bl

Figure 7-14: Scenario 7 itemized treatment costs for 243 mg/L sodium removed (Refer to Table

7-1 for full Scenario description)
130

Table 7-9: Scenario 7 Economic Parameter Percent Contribution to Total Treatment Cost for

243 mg/L Sodium Removed (Refer to Table 7-1 for full Scenario description).

Parameter IX HCl IX H2SO4 EDR HEEPM RO

Capital 6% 9% 47% 18% 9%

Brine 32% 45% 20% 15% 41%

Electric 0.1% 0.1% 6.4% 24.2% 6.3%

Labor 1% 2% 17% 8% 6%

Chemical 56% 39% 1% 26% 24%

Parts 4% 5% 4% 7% 11%

Iron Pretreatment 0% 0% 5% 2% 2%

Total 100% 100% 100% 100% 100%

Total Cost ($/bbl) 0.104 0.074 0.102 0.138 0.149

7.4 Summary

The total treatment costs for all technologies from all scenarios at 243 mg/L of sodium

removed are summarized in Table 7-10. Based on toolbox results, when treating PRB CBM

waters the cost effectiveness of treatment technologies are influenced by optional user inputs.

However, IX H2SO4 is the most cost effective technology up to the crossover point where EDR

becomes the most cost effective (approximately 410 to 760 mg/L of sodium removed depending

on the conditions). There are, however, safety concerns associated with IX H2SO4 which may

make it necessary to select the next most cost effective technology when removing lower

amounts of sodium, IX HCl or EDR.

Table 7-10: Total Treatment Costs for all scenarios at 243 mg/L of Sodium Removed. Refer to

Table 7-1 for full scenario descriptions. Units are $/bbl of influent water.

Scenario

Technology

1 2 3 4 5 6 7

IX HCl 0.111 0.104 0.113 0.078 0.105 0.081 0.104

IX H2SO4 0.069 0.063 0.071 0.036 0.063 0.063 0.074

EDR 0.149 0.102 0.132 0.085 0.115 0.102 0.102

HEEPM 0.148 0.124 0.145 0.107 0.190 0.124 0.138

RO 0.147 0.135 0.143 0.085 0.153 0.135 0.149
131

When utilizing the DTC toolbox note that the cost output should not be taken to be exact

treatment costs, but rather predictions of treatment costs based on published values and estimates

provided to the author. Costs presented are for a facility owned and operated by a CBM water

producer not another entity.

Also, extra costs such as, pipelines to get the water to and from the facility, roads,

equalization storage, water degassing, and profit to a consulting firm or treatment technology

provider have not been accounted for. These costs have not been included due to their variability

and the lack of data resulting from their proprietary nature. Therefore, it is strongly

recommended that the producer get detailed treatment cost estimates prior to deciding what

technology to utilize.
132

8. SUMMARY, CONCLUSIONS, AND RECOMMENDATIONS

8.1 Summary

The CBM-DTC Toolbox has been developed to aid industry, regulators and researchers

in deciding which demineralization technology to use and under what conditions when treating

PRB CBM produced waters. The first step in development was to define the constituent

concentrations of PRB CBM produced waters and use requirements. Then an examination of

currently used technologies, lead to the conclusion that demineralization technologies must be

used to remove sodium from these waters. Literature review and contact of technology vendors

then identified the most probable technologies that may be used to treat these waters: ion

exchange with hydrochloric acid or sulfuric acid regeneration, reverse osmosis, electrodialysis

reversal, and high efficiency electro-pressure membranes.

After determining the technologies that are most likely to be used to treat PRB CBM

waters, a general treatment flow schematic was developed and necessary mass balances were

performed to define functions for approximating flow rates and concentrations for the bypass,

treated, brine, and final streams. Functions were also developed to approximate the water

recovery and chemical rejection efficiency of selected technologies from the literature.

Capital, operating and maintenance costs were developed for each technology from the

literature. Cost algorithms developed took into account the influent water characteristics, effluent

characteristics, pretreatment requirements, and flow rates. Costs were also developed for the

following brine management technologies: deep well injection, evaporation ponds, and
133

evaporation with crystallization. The toolbox was constructed in a Microsoft Excel environment

with calculations being performed by underlying Visual Basic Macros.

Sensitivity analyses were then performed using the toolbox to determine the sensitivity of

treatment cost to cost factors such as: facility life, influent flow rate, sodium removed, chemical

cost, electrical cost, parts replacement costs, and brine disposal costs. Design Expert software

was used to perform these analyses.

The final task was to determine the impact of the above factors on technology cost

effectiveness. This analysis showed under what conditions one technology may be more cost

effective than another. In all scenarios tested ion exchange with sulfuric acid regeneration was

the most cost effective for lower amounts of sodium removed, and electrodialysis reversal was

the most cost effective for higher sodium removals.

8.2 Conclusions

The CBM-DTC Toolbox makes it possible to compare five different demineralization

technologies on an equal platform to determine preliminary treatment cost approximations, and

facilitate initial technology selection. Results show that PRB CBM water treatment costs may

range from $0.036/bbl to $0.190/bbl depending on the technology selected and the conditions

(Table 7-10).

Sensitivity analyses have shown that treatment cost is directly impacted by the amount of

sodium removed from the water (as the amount of sodium removed increases the treatment cost

increases regardless of the technology selected). Also, if the limiting constituent for the process

is not sodium the most cost effective technology may change. For example, if chloride or sulfate

must be removed from the water to meet a discharge requirement, then cation exchange will not
134

be able to treat the water to meet the requirement because chloride or sulfate can not be removed

from water via cation exchange. Sensitivity analyses also indicated that a treatment facility’s

flow rate and operational life should be maximized to decrease overall treatment cost.

Toolbox results have indicated that cation exchange is the most cost effective technology

for treating PRB CBM waters when low amounts of sodium must be removed from the water.

However, this result is attributed solely to the major anion in PRB CBM waters being

bicarbonate (HCO3

-

). If sulfate or chloride, for example, were the major anion(s) in PRB CBM

waters another technology, such as EDR, may become the most cost effective. This is attributed

to both anion and cation exchange being required which will increase the treatment cost for ion

exchange.

8.3 Recommendations

Based on the CBM-DTC Toolbox results a feasibility study of electrodialysis reversal

treatment should be performed in the PRB through the use of a pilot scale EDR treatment plant.

This may become more necessary in the future if brine disposal costs increase, which will make

EDR treatment even more competitive versus ion exchange treatment technologies.

PRB CBM producers and vendors should also work toward decreasing brine management

costs through all possible means including volume reduction and by-product generation/

utilization. Technologies that will aid in this task may include bipolar electrodialysis, membrane

distillation, and evaporation with crystallization. Most notably, sodium sulfate or sodium

chloride brines generated in ion exchange processes may be converted into viable products such

as sulfuric acid or hydrochloric acid through the use of bipolar ED.
135

The final recommendation is for PRB producers and treatment vendors to consider using

renewable resources such as, solar, wind, and water (energy recovery turbines) to aid in

powering water treatment operations. In theory after the initial capital investment these

technologies may reduce overall treatment costs considerably, especially for energy intensive

processes like reverse osmosis or electrodialysis reversal.
136

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Western Resources Project Final Report. 2005. Produced Groundwater Associated with Coalbed

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144

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145

10. APPENDICES

APPENDIX A: Treatment Technology Data and Analysis

Appendix A-P: Pretreatment

Determination of greensand pretreatment capital costs:

Capital costs included in this analysis are illustrated in Table AP-1. These costs include:

greensand filtration media, filtering vessels, backwash holding tanks, structure costs and

miscellaneous equipment costs. The overall capital cost is calculated using Equation AP.1.

CCpretreat = CCgreensand + CCvessel + CCstructure + CCtank + CCmisc.(AP.1)

Where, CCpretreat is the pretreatment capital cost for manganese greensand iron removal ($),

CCgreensand capital cost for greensand ($), CCvessel is the vessel capital cost ($), CCstructure is the

capital cost for a structure to house the system ($), CCtank is the capital cost for a backwash

holding tank ($), and CCmisc. is the capital cost of miscellaneous equipment ($).

Table AP-1: Greensand capital cost estimate from Hoffman et al. (2006) for required equipment

for iron removal with manganese greensand. Miscellaneous refers to other equipment costs such

as: feed pumps, piping, mixing tanks, and blowers.

Item Cost

Greensand Filter Media $90 per ft

3

Vessel $125,000 per 9ft diameter

Backwash Holding Tank $30,000 per 127 ft

2

Structure $100 per ft

2

Miscellaneous $19,000

The capital cost for greensand was determined first. This was done by determining the

required filter area for different feed water flow rates at a filter loading rate equal to 4 gpm/ft

2

(Hoffman et al., 2006). The required greensand volume was then found by multiplying the filter
146

area by the bed depth, 3ft (Inversand Company, 2006). Finally, the greensand cost was found by

multiplying the media volume by the unit cost of manganese greensand from Table AP-1.

Next, the required vessel diameter was found from the calculated area. This diameter

was used to calculate the cost for filtering (pressure) vessels. The cost of a pressure vessel is

assumed to be directly proportional to its diameter. Pressure vessel cost is calculated with

Equation AP.2.













= ⋅

9ft

$125,000

CCvessel Dvessel

(AP.2)

Where, $125,000 is the cost for a 9ft diameter pressure vessel from Table AP-1.

The cost of a structure to house the system was next determined. The space required for

iron removal with greensand was assumed to be equal to the required filter area plus 20% of the

filtering area (Young, 2006). The unit cost for the structure was then assumed to be equal to the

value from Table AP-1 ($100/ft

2

). The structure capital cost is calculated with Equation AP.3.

2

structure filter CC =1.2 ⋅ A ⋅ $100/ft (AP.3)

Where, 1.2·Afilter is the filter area plus 20% and $100/ft

2

is the unit cost of the structure.

The cost for a backwash holding tank is also dependent on the required filter area. This

cost was adapted directly from Hoffman et al. (2006) in the same manner as vessel capital costs.

The capital cost for a backwash holding tank is calculated with Equation AP.4.













= ⋅

tank filter 2

127ft

$30,000

CC A (AP.4)

Where, $30,000 is the cost for a backwash tank capable of holding two backwashes for a system

with a 127ft

2

filter area from Table AP-1.

Miscellaneous costs which include feed pumps, piping, mixing tanks, and blowers are

considered constant with respect to Qp and filter area. These costs are considered constant
147

because they are small with respect to other costs involved in this process, (see Table AP-1).

Miscellaneous costs are equal to $19,000.

All the costs discussed above were combined in a spreadsheet and a relationship was

found that relates Qp to iron pretreatment capital costs, Figure 3-3. Equation 3.11 calculates the

capital cost for iron removal from water by greensand filtration.

Appendix A-RO: Reverse Osmosis

y = -0.0015x + 100

R

2

= 1

0

10

20

30

40

50

60

70

80

90

100

0 5000 10000 15000 20000 25000 30000 35000 40000

Feed TDS, mg/L

W

ater Reco

very, %

Figure ARO-1: Reverse osmosis water recovery with respect to influent TDS. Influent feed

pressure equals 800psi for predominately NaCl feed solution while maintaining a constant

effluent TDS concentration equal to 500mg/L. Source: EET Corporation, 3106 Roane State

Hwy, Harriman, TN 37748 www.eetcorp.com.
148

0.000

0.050

0.100

0.150

0.200

0.250

0.300

0.350

0 1000 2000 3000 4000 5000 6000 7000 8000

Feed TDS, mg/L

Electrical Usa

ge, k

W-hr/b

bl

Figure ARO-2: Electrical energy requirement with respect to feed TDS for reverse osmosis.

Adapted from Leitner and Associates (1997).

Table ARO-1: Capital Costs for RO with data source and ENR cost indexes in the year of

analysis and January 2007.

Source: Flow, mgd Capital Cost ENR Yr ENR 2007 2007 cost

0.0025 $41,600 6095 7879.6 $53,780

0.01 $57,200 6095 7879.6 $73,948

0.05 $112,900 6095 7879.6 $145,957

0.10 $187,500 6095 7879.6 $242,400

0.50 $628,600 6095 7879.6 $812,652

International

Consultants et al.,

2000

(Water Model)

1.00 $1,157,600 6095 7879.6 $1,496,542

1.00 $775,820 6095 7879.6 $1,002,978

10.0 $5,467,390 6095 7879.6 $7,068,227

100 $43,161,740 6095 7879.6 $55,799,384

International

Consultants et al.,

2000

(WATERCO$T

Model) 200 $80,961,170 6095 7879.6 $104,666,388

0.25 $335,000 5408 7879.6 $488,104

0.50 $605,000 5408 7879.6 $881,501

Beardsley et al.,

1994

1.00 $1,110,000 5408 7879.6 $1,617,300

Average 1.00 $1,372,273
149

Determination of Electrical use using the Taylor and Jacobs (1996) method

To develop a function that relates RO electrical use to feed TDS the following procedure

was used:

1) Brine and permeate TDS concentrations were found for several TDS concentrations:

a) Permeate TDS concentrations were found by multiplying the feed TDS

concentration by one minus the TDS rejection efficiency, 0.99.

b) Brine TDS concentrations were found by performing a mass balance on TDS,

Equations ARO.1 and ARO.2.

( ) [ ( )] f,TDS

-5 Qb = Qf

⋅ 1− Water Recovery = Qf⋅ 1− 1-1.5×10 ×C (ARO.1)

( ) ( )

b

f f,TDS t,TDS f b

b,TDS Q

Q C C Q Q

C

⋅ − ⋅ −

= (ARO.2)

Where, Qb and Qf are the brine and feed flow rates, respectively (bbl/day), and

Cf,TDS, Ct,TDS and Cb,TDS are the feed, treated and brine TDS concentrations,

respectively (mg/L)

2) The average osmotic pressure gradient was found for several feed TDS concentrations

using equation ARO.3 from Taylor and Jacobs (1996).

C 0.01

2

C C

	 t,TDS

f,TDS b,TDS

sys

⋅

















−

+

= (ARO.3)

Where, sys is the average osmotic pressure gradient for the system (lb/in

2

), Cf,TDS is the

feed TDS (mg/L), Cb,TDS is the concentrate or brine TDS concentration (mg/L), Ct,TDS is

the permeate or treated water TDS concentration (mg/L). and 0.01 is an approximation

factor for converting TDS (mg/L) to pressure (lb/in

2

) which is analogous to RT in

Equation 3.20.
150

3) The required feed pressure was found using equation ARO.4 with the assumptions listed

in Table ARO-2 from Taylor and Jacobs (1996):

( ) = 

	











 − − ∆









 − × − ×

= × Γ

n

i

p i

F

w i P

P i L

J K

1

( ) ( )

2

2 2 1

π (ARO.4)

4) The Energy requirement was found using equation ARO.5. See Table ARO-3 for results

from Equations ARO.1, ARO.2, ARO.3, ARO.4 and ARO.5

( )

r p m

3

2

2

F

  

0.00315

62.4 lb

1ft

ft

in

P 144

E

⋅ ⋅

































⋅

= (ARO.5)

Where, E is the energy required (kWh/kgal), PF is the feed pressure (lb/in

2

), r is the

water recovery (calculated with Equation 3.23), p is the pump efficiency (0.85), m is the

motor efficiency (0.85), 1ft

3

/62.4lb is the water density, and 0.00315 is a conversion

factor used to convert feet of water head to kWh/1000gal.

Table ARO-2: Variable definitions and values assumed for equation ARO.4

Variable Definition Value assumed unit

J Design Water Flux 15

1

gal/ft

2

-day

KW Membrane mass transfer Coefficient for water 0.3

1

gal/ft

2

-day /lb/in

2

i Stage number


(i) Array flow weighted factor for stage i 1

1

Pel Pre-element pressure loss 3

1

lb/in

2

Pe Stage entrance pressure loss 5

1

lb/in

2

Nel Number of elements per pressure vessel 6

2

n Number of stages 2

3

L Sum of mech. head losses = (Pel · Nel) + (Pe · n) 28

4

lb/in

2

	(i) Osmotic pressure gradient for stage i 	sys lb/in

2

Pp Desired Permeate Pressure 30

1

lb/in

2

PF Feed Pressure Calculated lb/in

2

1

value adapted from Taylor and Jacobs (1996).

2

value from IMSDesign.

3User defined.

4Calculated with user defined values and values from other sources.
151

Table ARO-3: Values Used and Calculated to Find a Function Relating Electrical Usage to Feed

TDS for RO (Equation 3.27 and Figure 3-9). See Equations ARO.1, ARO.2, ARO.3, ARO.4,

ARO.5, and Table ARO-2 for variable definitions. PF was found using Microsoft Excel Solver.

Feed TDS,

mg/L

Recovery Brine TDS,

mg/L

Effluent TDS,

mg/L

 sys,

psi

PF,

psi

Head,

ft

Energy,

kWh/1000gal

Energy,

kWh/bbl

100 0.9985 66334 1 332 423 977 4.2639 0.1791

1500 0.9775 66341 15 339 430 992 4.4265 0.1859

6000 0.91 66363 60 361 452 1044 4.9998 0.2100

10000 0.85 66383 100 381 472 1089 5.5859 0.2346

20000 0.7 66433 200 430 521 1203 7.4908 0.3146

35000 0.475 66508 350 504 595 1373 12.6039 0.5294

Appendix A-IX: Ion Exchange

Table AIX-1: Ion exchange capital costs adjusted with ENR cost Indexes. Adapted from Rubel

(2003).

Flow rate, gpm Vresin, ft

3 Capital, 2003 ENR 2003 ENR 2007 Capital, 2007

65 79 $332,000 6694.64 7879.6 $390,764

115 139 $377,000 6694.64 7879.6 $443,729

230 278 $472,000 6694.64 7879.6 $555,545

330 399 $540,000 6694.64 7879.6 $635,581

480 581 $602,000 6694.64 7879.6 $708,555

555 671 $705,000 6694.64 7879.6 $829,786

620 750 $736,000 6694.64 7879.6 $866,273

700 847 $761,000 6694.64 7879.6 $895,698
152

Appendix A-ED: Electrodialysis

y = 0.0002x

2 + 90.36x + 1E+06

R2 = 0.9476

0

5,000,000

10,000,000

15,000,000

20,000,000

25,000,000

30,000,000

0 50,000 100,000 150,000 200,000 250,000

Capital Cost, $

Feed Flow Rate, bbl/day

Leitner and Associates, 1997

Urano, 1977

Larson and Leitner, 1979

Combined

Figure AED-1: EDR capital costs as a function of feed water flow rate (capacity) with data

illustrated by source. Combined refers to the data set consisting of all data sources. Urano, 1997

(1000mg/L) is data developed by Urano (1977) for an EDR system designed for 1,000mg/L TDS

removal, and Larson and Leitner, 1979 (1300mg/L) is data developed by Larson and Leitner

(1979) for an EDR system designed for 1300mg/L TDS removal.

y = 150.8x + 1E+06

R2 = 0.996

0

5,000,000

10,000,000

15,000,000

20,000,000

25,000,000

30,000,000

35,000,000

40,000,000

0 50,000 100,000 150,000 200,000 250,000

Capital Cost, $

Feed Flow Rate, bbl/day

Urano, 1977

Leitner and Associates, 1997

Larson and Leitner, 1979

Lawrence et al., 1998

Combined

Figure AED-2: EDR capital costs as a function of feed water flow rate (capacity) with data

illustrated by source. Combined refers to the data set consisting of all data sources. Urano, 1997

(3000mg/L) is data developed by Urano (1977) for an EDR system designed for 3,000mg/L TDS

removal, and Larson and Leitner, 1979 (1300mg/L) is data developed by Larson and Leitner

(1979) for an EDR system designed for 1300mg/L TDS removal.

.
153

y = 355.9x + 1E+06

R2 = 0.992

0

5,000,000

10,000,000

15,000,000

20,000,000

25,000,000

0 10,000 20,000 30,000 40,000 50,000 60,000 70,000

Capital Cost, $

Feed Flow Rate, bbl/day

Urano, 1977

Lawrence et al., 1998

Combined

Figure AED-3: EDR capital costs as a function of feed water flow rate (capacity) with data

illustrated by source. Combined refers to the data set consisting of all data sources. Urano, 1997

(10000mg/L) is data developed by Urano (1977) for an EDR system designed for 10,000mg/L

TDS removal.

Table AED-1: Leitner and Associates (1997) EDR capital cost data adjusted to January 2007

dollars with ENR Indexes. Flow is the system capacity (feed water flow rate) and ENR Yr built

is the ENR index in the year built.

Flow, gpd Capital Yr built Yr built ENR Yr built 2007 Capital

100,000 $387,942 1975 2212 $1,381,929

300,000 $1,200,000 1984 4146 $2,280,637

432,000 $900,000 1995 5471 $1,296,224

500,000 $1,135,700 1990 4732 $1,891,137

1,000,000 $1,050,000 1989 4615 $1,792,758

1,900,000 $1,492,600 1993 5210 $2,257,407

3,500,000 $6,100,000 1989 4615 $10,415,073

3,750,000 $6,920,000 1990 4732 $11,522,999

4,500,000 $5,000,000 1993 5210 $7,561,996

Table AED-2: Lawrence et al. (1998) EDR capital cost for various TDS removals for a system

capacity of 168,000gpd adjusted to January 2007 dollars.

TDS Removed 5000mg/L 7500mg/L 9000mg/L

1998 Capital $1,044,000 $1,387,000 $1,575,000

2007 Capital $1,389,578 $1,846,116 $2,096,346
154

Table AED-3: Urano (1977) EDR capital cost for various flow rates and TDS removals adjusted

to January 2007 dollars. Flow is the system capacity (feed water flow rate).

Flow, gpd Capital Cost with TDS Removed 2007 Capital Cost with TDS Removed

1000mg/L 3000mg/L 10000mg/L 1000mg/L 3000mg/L 10000mg/L

26,400 $300,000 $375,000 $425,000 $917,655 $1,147,069 $1,300,012

52,800 $325,000 $400,000 $500,000 $994,127 $1,223,540 $1,529,425

79,200 $400,000 $500,000 $600,000 $1,223,540 $1,529,425 $1,835,311

132,000 $475,000 $525,000 $800,000 $1,452,954 $1,605,897 $2,447,081

264,000 $550,000 $800,000 $1,125,000 $1,682,368 $2,447,081 $3,441,207

528,000 $825,000 $1,125,000 $2,000,000 $2,523,552 $3,441,207 $6,117,702

792,000 $1,125,000 $1,750,000 $3,000,000 $3,441,207 $5,352,989 $9,176,553

1,320,000 $1,500,000 $2,050,000 $4,250,000 $4,588,276 $6,270,644 $13,000,116

2,640,000 $2,000,000 $4,000,000 $7,500,000 $6,117,702 $12,235,404 $22,941,382

Table AED-4: Larson and Leitner (1979) EDR capital cost for various flow rates and TDS

removals adjusted to January 2007 dollars. Flow is the system capacity (feed water flow rate).

Flow, gpd Capital w/ TDS Removed 2007 Capital w/ TDS Removed

1300mg/L 3148mg/L 1300mg/L 3148mg/L

1,000,000 $1,300,000 $1,700,000 $3,411,082 $4,460,646

3,000,000 $3,600,000 $4,650,000 $9,446,074 $12,201,179

5,000,000 $5,500,000 $7,500,000 $14,431,502 $19,679,321

10,000,000 $10,000,000 $14,000,000 $26,239,094 $36,734,732
155

Appendix A-H: HEEPMTM

Table AH-1: HEEPMTM capital cost estimates and design parameters for two system capacities,

15,000 and 45,000 bbl/day (EET, 2006).

Item

HEEPMTM HEEPMTM

System Capacity Rate (bbl/day) 15,000 45,000

System Capacity Rate (gpm) 446.8 1,340.4

System Capacity Rate (gpd) 630,000 1,890,000

System Feed (approx. ppm, as TDS) 4,400 4,400

Product Quality (approx. ppm, as TDS) 250 250

Brine Concentration (ppm, as TDS) 200,000 200,000

Brine Production Rate (BWPD) 311 934

Brine Production Rate (gpd) 13,073 39,218

Total System Recovery (%) 97.9% 97.9%

Hours of System Operation (hrs/day) 23.5 23.5

HEED® Model No. 400-B 400-B

No. Cells 1,100 3,200

Work Tank Size 10,000 2@10,000

Number RO Modules/Vessels 70/35 210/105

Total Energy Requirement (kw-hr/1000 gal) 27.5 27.5

Shelter Space Requirement (approx. ft

2

) 2500 3000

Total System Cost (US Dollars) $2,680,000` $4,990,000

Table AH-2: HEEPMTM itemized operating and maintenance costs for a PRB CBM produced

water with a TDS concentration of 4400 mg/L and assumptions (EET, 2006).

Item $/1000gal Assumptions

Electricity $2.20 27.5 kW-hr/ 1000gal at $0.08/kW-hr

Labor $0.36 $50,000 /employee-yr

Sulfuric Acid $1.18 2.2gal 98% H2SO4/1000gal water at $70/ton

Membrane Replacement $0.30 RO = 3yrs ED = 5yrs
156

Figure AH-1: HEEPMTM capital cost versus system capacity and feed TDS (EET, 2006).

Q<238

y = 371.48x + 75096

R

2

= 0.9524

$80,000

$100,000

$120,000

$140,000

$160,000

$180,000

$200,000

0 50 100 150 200 250

Feed Water Flow Rate (bbl/day)

S

ystem

C

a

pital C

ost

Figure AH-2: HEEPMTM capital costs as a function of feed water flow rates less than or equal to

238 bbl/day
157

238<Q<2381

y = 208.29x + 144883

R

2

= 0.9933

$50,000

$100,000

$150,000

$200,000

$250,000

$300,000

$350,000

$400,000

$450,000

$500,000

0 500 1,000 1,500 2,000 2,500

Feed Water Flow Rate (bbl/day)

S

ystem

C

a

pital C

ost

Figure AH-3: HEEPMTM capital costs as a function of feed water flow rates greater than 238

bbl/day and less than or equal to 2,381 bbl/day

2381<Q<23810

y = 160.99x + 259932

R

2

= 0.9979

$500,000

$1,000,000

$1,500,000

$2,000,000

$2,500,000

$3,000,000

2,500 5,000 7,500 10,000 12,500 15,000 17,500 20,000 22,500 25,000

Feed Water Flow Rate (bbl/day)

S

ystem

C

a

pital C

ost

CC_adjusted

Estimate pts

Figure AH-4: HEEPMTM capital costs as a function of feed water flow rates greater than 2,381

bbl/day and less than or equal to 23,810 bbl/day
158

23810<Q

y = 111.47x

R

2

= 1

$4,000,000

$5,000,000

$6,000,000

$7,000,000

$8,000,000

$9,000,000

$10,000,000

$11,000,000

40,000 60,000 80,000 100,000

Feed Water Flow Rate (bbl/day)

S

ystem

C

a

pital C

ost

CC_adjusted

Estimate pts

Figure AH-5: HEEPMTM capital costs as a function of feed water flow rates greater than 23,810

bbl/day

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

1,000 10,000 100,000 1,000,000

Feed Water Flow Rate (bbl/day)

35000

20000

10000

5000

1500

Figure AH-6: HEEPMTM operating and maintenance costs as a function of feed TDS and flow

rate with an effluent TDS equal to 500 mg/L. (EET, 2006).

Explanation of Electrical Use estimation for HEEPM

EET Corp. provided electrical use data for a system treating 4400 mg/L feed TDS from

the estimate Table AH-2. The electrical use was 1.16 kW-hr/bbl. EET Corp also provided Figure
159

AH-6 which relates operating and maintenance costs to feed TDS and feed water flow rate.

These sources were used in the following way to determine electrical and other costs.

1) The total operating and maintenance cost was found from the estimate, and Figure AH-6

for a feed water flow rate equal to 15000 bbl/day and feed TDS equal to 4400 mg/L.

The resulting annual operating and maintenance costs were $500,000/yr from Figure AH6 and $896,216/yr from the estimate.

2) A ratio ( ratio 1) was found between these numbers equal to 1.79.

3) The electrical cost per year was determined from the estimate to be $505,890/yr at a unit

cost of $0.08/kW-hr.

4) A ratio (ratio 2) was found between electrical cost and total operating and maintenance

cost from the estimate ($896,216/yr) equal to 0.56

5) The actual operating and maintenance cost for each TDS regime illustrated in Figure AH6 were adjusted with ratio 1 to give the actual operating and maintenance cost for PRB

CBM waters.

6) The electrical cost was found by multiplying the actual operating and maintenance cost

by ratio 2.

7) The electrical cost was divided by the unit cost of electricity $0.08/kW-hr.

8) These steps were repeated for several flows and all TDS regimes illustrated in Figure

AH-6 to develop a function relating electrical usage to feed TDS concentration Figure 3-

26 and Equation 3.63.

Explanation of Alkalinity used to determine Sulfuric acid dose for HEEPM

The estimate provided by EET Corp. was for a water with a feed TDS concentration

equal to 4400 mg/L. The alkalinity of this water was not provided. To determine the alkalinity of
160

this water based on feed TDS alone the following was done.

1) The ratio of alkalinity (2320 mg/L) to feed TDS (2010 mg/L) for Powder River High in

Table 2-1 was found.

2) The feed TDS for the water in the estimate was multiplied by the ratio (1.154)

3) The resulting value is the assumed alkalinity of the water in the estimate (5078 mg/L as

CaCO3) .
161

APPENDIX B: BRINE MANAGEMENT ANALYSIS

Energy Required to Evaporator water with and without an evaporator

Evaporation w/o an evaporator:

Energy required to evaporate 1L of water from an initial temperature equal to 19 oC:

Energy to 100 oC = (mass water, 1 kg) (specific heat, 4.187 kJ kg

-1

oC

-1

) (100 oC - 19 oC)

Energy to 100 oC = 339 kJ

Energy to boil water to steam = (heat of vaporization, 2270 kJ/kg) (mass water, 1 kg)

Energy to boil water to steam = 2270 kJ

Total Energy Required = 339 kJ + 2270 kJ = 2609 kJ = 0.725 kW-hr/L distillate

Evaporation w/ an evaporator:

L of brine

kW hr

0.033

3.78 L

1gal

42 gal

1bbl

bbl of brine

kW hr

Energy 5.26

−

⋅ ⋅ =

−

=

Concentrate from 2% to 20% solids by weight

Assume: brine density = 1000 g/L

2%  980 g water and 20 g solids

20%  80 g water and 20 g solids

Water Removed from 1 L of brine as distillate = (980 g – 80 g) (1L / 1000 g) = 0.9 L

L of Distillate

kW hr

0.0367

0.9 L of Distillate

1L of brine

L of brine

kW hr

Energy 0.033

−

⋅ =

−

=
162

APPENDIX C: LIST OF REVIEWED WYPDES PERMITS

Active WYPDES permits reviewed:

0043958, 0044300, 0051934, 0050270, 0053945, 0042412, 0041611, 0046311, 0037699,

0039641, 0047597.

Watershed based permits reviewed:

Pumpkin Creek, Four Mile Creek, Willow Creek, Fence Creek, Clear Creek.

Watershed based permits in permitting process reviewed:

Lower Tongue River, Crazy Woman Creek, Deadhorse Creek.
163

APPENDIX D: Factorial design experimental setups and toolbox results

Table D-1: Reverse osmosis factorial design with responses

Run

Acid Cost,

$/ton

Membrane

Cost, $/bbl

Facility Life,

yr

Brine Cost,

$/bbl

Na Removed,

mg/L

Flow,

bbl/day

Electric,

$/kW-hr

Response

$/bbl

1 250 0.034 10 1 500 100000 0.07 0.092

2 250 0.034 10 5 500 100000 0.21 0.17

3 135 0.017 5 5 500 10000 0.21 0.201

4 250 0.017 5 5 500 100000 0.21 0.194

5 250 0.017 10 5 500 10000 0.21 0.19

6 250 0.034 5 1 500 100000 0.21 0.128

7 250 0.034 10 5 100 10000 0.21 0.04

8 135 0.017 10 5 100 10000 0.21 0.039

9 250 0.034 5 1 500 10000 0.07 0.126

10 250 0.034 10 1 100 10000 0.07 0.024

11 250 0.017 10 5 500 100000 0.07 0.158

12 135 0.034 10 5 500 100000 0.07 0.134

13 135 0.034 5 5 100 10000 0.21 0.046

14 135 0.017 5 1 500 10000 0.07 0.123

15 135 0.017 5 5 500 100000 0.07 0.158

16 250 0.017 5 1 500 10000 0.21 0.158

17 250 0.034 5 5 500 10000 0.21 0.205

18 250 0.017 10 1 100 10000 0.21 0.031

19 135 0.017 5 1 100 10000 0.21 0.037

20 250 0.034 10 5 100 100000 0.07 0.031

21 250 0.017 5 5 100 100000 0.07 0.038

22 135 0.017 10 5 500 10000 0.07 0.155

23 250 0.034 10 1 100 100000 0.21 0.024

24 250 0.017 10 1 500 10000 0.07 0.112

25 135 0.034 10 5 100 100000 0.21 0.033

26 135 0.017 5 5 100 100000 0.21 0.039

27 135 0.034 5 1 100 100000 0.21 0.026

28 135 0.034 10 1 500 100000 0.21 0.101

29 135 0.017 10 5 100 100000 0.07 0.031

30 250 0.034 5 1 100 100000 0.07 0.024

31 250 0.034 5 5 100 10000 0.07 0.044

32 250 0.017 10 1 500 100000 0.21 0.124

33 250 0.017 5 5 100 10000 0.21 0.051

34 135 0.034 5 1 500 100000 0.07 0.092

35 135 0.034 10 5 500 10000 0.21 0.167

36 135 0.017 10 1 100 100000 0.21 0.024

37 135 0.034 5 1 100 10000 0.07 0.03

38 135 0.017 10 1 500 100000 0.07 0.088

39 135 0.034 5 1 500 10000 0.21 0.135

40 135 0.017 10 1 100 10000 0.07 0.024

41 135 0.017 5 1 500 100000 0.21 0.124

42 135 0.017 5 1 100 100000 0.07 0.023

43 135 0.017 5 5 100 10000 0.07 0.043

44 250 0.017 10 1 100 100000 0.07 0.022

45 250 0.017 5 5 500 10000 0.07 0.168

46 135 0.017 10 5 500 100000 0.21 0.166

47 250 0.017 5 1 100 10000 0.07 0.035

48 250 0.017 5 1 500 100000 0.07 0.115

49 250 0.017 10 5 100 10000 0.07 0.038

50 250 0.034 5 1 100 10000 0.21 0.037

51 250 0.034 10 5 500 10000 0.07 0.158

52 250 0.034 5 5 100 100000 0.21 0.04

53 250 0.034 5 5 500 100000 0.07 0.162

54 135 0.034 10 1 500 10000 0.07 0.089

55 135 0.034 10 1 100 10000 0.21 0.026

56 135 0.034 5 5 500 10000 0.07 0.169

57 135 0.017 10 1 500 10000 0.21 0.121

58 135 0.034 5 5 100 100000 0.07 0.033

59 135 0.034 10 5 100 10000 0.07 0.033

60 135 0.034 5 5 500 100000 0.21 0.17

61 250 0.017 10 5 100 100000 0.21 0.038

62 135 0.034 10 1 100 100000 0.07 0.017

63 250 0.034 10 1 500 10000 0.21 0.124

64 250 0.017 5 1 100 100000 0.21 0.031
164

Table D-2: Ion Exchange factorial design with responses

Run Acid Cost, $/ton Resin Cost, $/bbl

Facility Life,

yr

Brine Cost, $/bbl Na removed, mg/L

Flow,

bbl/day

Electric, $/kWhrResponse

$/bbl

1 135 0.0106 5 1 500 100000 0.07 0.102

2 135 0.0053 10 5 100 10000 0.21 0.038

3 250 0.0053 10 5 100 100000 0.21 0.04

4 135 0.0053 10 5 500 100000 0.21 0.141

5 135 0.0053 10 5 100 100000 0.07 0.029

6 135 0.0106 10 5 100 100000 0.21 0.029

7 135 0.0053 10 1 100 10000 0.07 0.028

8 250 0.0053 10 1 500 10000 0.07 0.155

9 250 0.0053 5 1 100 100000 0.21 0.033

10 135 0.0106 5 5 100 100000 0.07 0.032

11 135 0.0106 10 1 500 100000 0.21 0.091

12 250 0.0106 5 5 100 10000 0.07 0.06

13 250 0.0106 5 5 500 100000 0.07 0.203

14 250 0.0106 10 5 500 100000 0.21 0.192

15 250 0.0053 5 1 500 100000 0.07 0.157

16 250 0.0106 10 1 100 10000 0.07 0.039

17 135 0.0053 10 1 100 100000 0.21 0.02

18 135 0.0053 10 1 500 100000 0.07 0.093

19 250 0.0106 10 1 500 100000 0.07 0.144

20 135 0.0053 10 5 500 10000 0.07 0.149

21 135 0.0053 5 5 500 100000 0.07 0.152

22 250 0.0053 5 1 100 10000 0.07 0.051

23 250 0.0106 5 1 500 100000 0.21 0.155

24 135 0.0106 5 1 500 10000 0.21 0.119

25 135 0.0053 5 5 500 10000 0.21 0.169

26 135 0.0106 5 1 100 100000 0.21 0.022

27 250 0.0053 10 1 100 100000 0.07 0.03

28 135 0.0106 5 5 100 10000 0.21 0.049

29 135 0.0053 5 1 100 100000 0.07 0.023

30 250 0.0053 5 1 500 10000 0.21 0.175

31 250 0.0106 5 1 100 10000 0.21 0.05

32 135 0.0106 10 1 100 100000 0.07 0.019

33 250 0.0106 5 5 500 10000 0.21 0.221

34 250 0.0053 5 5 100 10000 0.21 0.06

35 135 0.0053 5 1 500 10000 0.07 0.121

36 135 0.0106 10 1 500 10000 0.07 0.099

37 135 0.0106 10 5 500 100000 0.07 0.138

38 250 0.0053 10 1 100 10000 0.21 0.039

39 135 0.0106 5 5 500 100000 0.21 0.15

40 135 0.0106 5 5 500 10000 0.07 0.167

41 135 0.0053 5 1 500 100000 0.21 0.104

42 250 0.0053 10 5 500 100000 0.07 0.194

43 250 0.0106 10 5 100 100000 0.07 0.04

44 250 0.0106 10 1 500 10000 0.21 0.153

45 250 0.0053 5 5 500 100000 0.21 0.205

46 250 0.0053 5 5 100 100000 0.07 0.043

47 135 0.0106 10 1 100 10000 0.21 0.028

48 250 0.0106 5 5 100 100000 0.21 0.043

49 250 0.0053 10 5 100 10000 0.07 0.049

50 250 0.0053 10 1 500 100000 0.21 0.147

51 250 0.0106 5 1 100 100000 0.07 0.033

52 135 0.0106 10 5 500 10000 0.21 0.147

53 135 0.0053 10 1 500 10000 0.21 0.102

54 135 0.0053 5 5 100 100000 0.21 0.032

55 250 0.0053 5 5 500 10000 0.07 0.223

56 135 0.0053 5 5 100 10000 0.07 0.05

57 250 0.0106 10 5 500 10000 0.07 0.201

58 135 0.0053 5 1 100 10000 0.21 0.04

59 135 0.0106 10 5 100 10000 0.07 0.037

60 250 0.0053 10 5 500 10000 0.21 0.203

61 250 0.0106 10 5 100 10000 0.21 0.048

62 250 0.0106 10 1 100 100000 0.21 0.03

63 250 0.0106 5 1 500 10000 0.07 0.173

64 135 0.0106 5 1 100 10000 0.07 0.04
165

Table D-3: Electrodialysis reversal factorial design with responses

Run

Electric,

$/kW-hr

Membrane,

$/bbl

Facility Life,

yr

Brine Cost,

$/bbl

Na removed,

mg/L

Flow,

bbl/day

Response

$/bbl

1 0.21 0.00504 5 5 100 100000 0.131

2 0.07 0.00252 5 5 500 10000 0.188

3 0.21 0.00252 10 1 100 10000 0.111

4 0.21 0.00252 5 1 500 10000 0.184

5 0.07 0.00504 5 1 500 10000 0.165

6 0.21 0.00504 10 1 100 100000 0.082

7 0.21 0.00252 10 5 100 100000 0.096

8 0.07 0.00504 5 5 500 100000 0.134

9 0.21 0.00252 5 5 100 10000 0.185

10 0.07 0.00504 10 1 100 10000 0.097

11 0.07 0.00504 10 5 500 10000 0.12

12 0.07 0.00252 5 1 500 100000 0.119

13 0.07 0.00252 10 5 500 100000 0.1

14 0.07 0.00252 10 1 100 100000 0.076

15 0.21 0.00504 5 1 500 100000 0.13

16 0.21 0.00504 10 1 500 10000 0.116

17 0.07 0.00252 10 1 500 10000 0.106

18 0.07 0.00252 5 5 100 100000 0.125

19 0.21 0.00252 10 5 500 10000 0.139

20 0.21 0.00504 5 1 100 10000 0.17

21 0.21 0.00252 5 5 500 100000 0.153

22 0.07 0.00504 10 1 500 100000 0.077

23 0.21 0.00252 10 1 500 100000 0.095

24 0.21 0.00504 10 5 100 10000 0.117

25 0.07 0.00504 5 5 100 10000 0.171

26 0.07 0.00252 5 1 100 10000 0.164

27 0.07 0.00504 5 1 100 100000 0.111

28 0.07 0.00504 10 5 100 100000 0.083

29 0.21 0.00252 5 1 100 100000 0.124

30 0.21 0.00504 10 5 500 100000 0.11

31 0.21 0.00504 5 5 500 10000 0.198

32 0.07 0.00252 10 5 100 10000 0.112
166

Table D-4: HEEPM factorial design with responses

Run Acid Cost,

$/ton

Membrane,

$/bbl

Facility Life,

yr

Brine Cost,

$/bbl

Na Removed,

mg/L

Flow,

bbl/day

Electric,

$/kW-hr

Response

$/bbl

1 135 0.0252 5 1 100 100000 0.21 0.052

2 135 0.0126 5 1 500 100000 0.21 0.231

3 250 0.0126 10 5 500 100000 0.07 0.164

4 135 0.0126 10 1 100 10000 0.07 0.037

5 250 0.0252 5 1 500 10000 0.07 0.216

6 135 0.0126 10 1 100 100000 0.21 0.044

7 250 0.0126 5 1 500 10000 0.21 0.299

8 250 0.0252 10 1 500 100000 0.07 0.14

9 135 0.0126 10 5 100 10000 0.21 0.056

10 135 0.0252 5 5 100 10000 0.21 0.073

11 250 0.0126 5 5 100 100000 0.07 0.046

12 250 0.0252 10 5 500 10000 0.07 0.182

13 135 0.0126 10 5 500 100000 0.21 0.22

14 250 0.0126 5 1 500 100000 0.07 0.174

15 250 0.0252 5 1 500 100000 0.21 0.247

16 135 0.0126 5 5 500 100000 0.07 0.172

17 250 0.0126 5 1 100 10000 0.07 0.059

18 135 0.0252 10 5 100 100000 0.21 0.047

19 250 0.0126 10 1 500 100000 0.21 0.223

20 135 0.0252 5 5 500 100000 0.21 0.244

21 135 0.0252 10 5 500 100000 0.07 0.137

22 135 0.0126 10 5 100 100000 0.07 0.143

23 250 0.0252 10 5 500 100000 0.21 0.236

24 135 0.0126 5 5 500 10000 0.21 0.296

25 135 0.0126 5 1 100 10000 0.21 0.07

26 250 0.0252 5 5 100 10000 0.07 0.061

27 250 0.0126 10 1 100 10000 0.21 0.057

28 250 0.0126 10 5 500 10000 0.21 0.265

29 135 0.0252 10 1 500 100000 0.21 0.197

30 135 0.0126 5 5 100 100000 0.21 0.057

31 250 0.0252 5 1 100 100000 0.07 0.041

32 135 0.0252 10 1 500 10000 0.07 0.142

33 135 0.0252 5 5 100 100000 0.07 0.041

34 135 0.0126 10 1 500 10000 0.21 0.225

35 250 0.0252 10 5 100 100000 0.07 0.035

36 135 0.0126 5 1 100 100000 0.07 0.038

37 250 0.0252 10 1 100 10000 0.07 0.04

38 135 0.0252 5 1 100 10000 0.07 0.053

39 135 0.0252 5 1 500 10000 0.21 0.272

40 135 0.0126 5 5 100 10000 0.07 0.058

41 250 0.0126 10 5 100 100000 0.21 0.052

42 250 0.0252 5 1 100 10000 0.21 0.073

43 135 0.0126 10 5 500 10000 0.07 0.166

44 250 0.0126 5 1 100 100000 0.21 0.058

45 250 0.0252 10 1 100 100000 0.21 0.047

46 250 0.0126 10 1 100 100000 0.07 0.033

47 135 0.0252 5 5 500 10000 0.07 0.213

48 250 0.0126 10 1 500 10000 0.07 0.145

49 250 0.0252 10 1 500 10000 0.21 0.241

50 135 0.0126 10 1 500 100000 0.07 0.124

51 250 0.0252 5 5 100 100000 0.21 0.06

52 135 0.0252 5 1 500 100000 0.07 0.148

53 250 0.0252 5 5 500 100000 0.07 0.187

54 250 0.0126 5 5 500 100000 0.21 0.27

55 135 0.0252 10 5 100 10000 0.07 0.039

56 135 0.0252 10 5 500 10000 0.21 0.238

57 135 0.0252 10 1 100 100000 0.07 0.027

58 250 0.0252 5 5 500 10000 0.21 0.312

59 250 0.0126 5 5 500 10000 0.07 0.239

60 135 0.0252 10 1 100 10000 0.21 0.051

61 250 0.0126 5 5 100 10000 0.21 0.078

62 250 0.0126 10 5 100 10000 0.07 0.045

63 135 0.0126 5 1 500 10000 0.07 0.2

64 250 0.0252 10 5 100 10000 0.21 0.059
167

Table D-5: Sensitivity analysis influent water chemistry

Constituent Concentration Units

Total dissolved solids 1074 mg/L

F

–

747 μg/L

Cl

–

10 mg/L

SO4

2–

3 mg/L

Br 46 μg/L

Alkalinity (as CaCO3) 674 mg/L

NH4

+

2 mg/L

Ca

2+

23 mg/L

K 6 mg/L

Mg

2+

12 mg/L

Na

+

600 mg/L

Ba

2+

441 μg/L

Fe 621 μg/L

Si 3 mg/L

Sr 523 μg/L

Al 6 μg/L

As 2 μg/L

Se 1 μg/L

B 51 μg/L

Cu 12 μg/L

Mo 1 μg/L

Mn 14 μg/L

Cr 43 μg/L

Zn 11 μg/L
168

Table D-6: Electrodialysis reversal factorial design with responses (labor included)

Run Electric Brine Facility Life Flow rate Sodium Removed Membrane Cost Labor Cost Response

$/kW-hr $/bbl brine yr bbl/day mg/L $/bbl $/bbl $/bbl

1 0.07 5 10 10000 100 0.00504 0.0096 0.104

2 0.21 1 10 10000 100 0.00504 0.0096 0.103

3 0.07 5 10 100000 100 0.00252 0.0096 0.075

4 0.21 1 10 10000 500 0.00252 0.0096 0.108

5 0.07 1 10 100000 500 0.00252 0.0096 0.069

6 0.21 1 10 100000 100 0.00504 0.0177 0.086

7 0.21 1 10 100000 100 0.00252 0.0096 0.074

8 0.07 5 10 100000 500 0.00504 0.0096 0.092

9 0.07 5 5 10000 100 0.00504 0.0177 0.175

10 0.07 1 5 100000 100 0.00504 0.0177 0.115

11 0.21 1 10 10000 100 0.00252 0.0177 0.107

12 0.07 5 10 100000 100 0.00504 0.0177 0.087

13 0.21 1 5 10000 500 0.00504 0.0096 0.176

14 0.21 5 5 100000 500 0.00504 0.0096 0.145

15 0.21 5 10 100000 500 0.00252 0.0096 0.102

16 0.07 1 10 100000 100 0.00252 0.0177 0.072

17 0.21 5 5 100000 100 0.00504 0.0177 0.135

18 0.21 5 5 10000 100 0.00252 0.0177 0.181

19 0.21 5 5 100000 100 0.00252 0.0096 0.123

20 0.21 5 5 10000 500 0.00504 0.0177 0.202

21 0.07 1 10 10000 100 0.00504 0.0177 0.101

22 0.07 1 10 10000 100 0.00252 0.0096 0.089

23 0.07 5 5 10000 100 0.00252 0.0096 0.163

24 0.07 5 5 100000 100 0.00252 0.0177 0.121

25 0.07 5 5 100000 500 0.00504 0.0177 0.138

26 0.07 5 10 10000 500 0.00504 0.0177 0.124

27 0.07 5 5 10000 500 0.00252 0.0177 0.184

28 0.07 5 5 10000 500 0.00504 0.0096 0.18

29 0.07 5 10 10000 100 0.00252 0.0177 0.108

30 0.07 1 5 100000 500 0.00252 0.0177 0.115

31 0.07 1 10 100000 100 0.00504 0.0096 0.068

32 0.21 5 10 10000 100 0.00252 0.0096 0.109

33 0.07 1 10 100000 500 0.00504 0.0177 0.081

34 0.21 5 5 10000 500 0.00252 0.0096 0.19

35 0.07 1 10 10000 500 0.00504 0.0096 0.098

36 0.21 5 5 100000 500 0.00252 0.0177 0.148

37 0.07 5 5 100000 500 0.00252 0.0096 0.13

38 0.21 1 5 100000 100 0.00252 0.0177 0.12

39 0.21 1 5 10000 100 0.00252 0.0096 0.162

40 0.21 1 5 100000 500 0.00504 0.0177 0.134

41 0.21 1 10 10000 500 0.00504 0.0177 0.12

42 0.07 5 10 100000 500 0.00252 0.0177 0.096

43 0.21 5 10 10000 100 0.00504 0.0177 0.121

44 0.21 5 5 10000 100 0.00504 0.0096 0.177

45 0.07 5 10 10000 500 0.00252 0.0096 0.112

46 0.21 1 5 10000 100 0.00504 0.0177 0.174

47 0.07 1 5 100000 500 0.00504 0.0096 0.111

48 0.07 1 10 10000 500 0.00252 0.0177 0.102

49 0.21 1 5 100000 100 0.00504 0.0096 0.116

50 0.07 1 5 10000 500 0.00504 0.0177 0.169

51 0.07 1 5 10000 500 0.00252 0.0096 0.157

52 0.07 1 5 10000 100 0.00504 0.0096 0.156

53 0.21 1 10 100000 500 0.00252 0.0177 0.091

54 0.21 1 5 100000 500 0.00252 0.0096 0.122

55 0.21 5 10 100000 100 0.00252 0.0177 0.092

56 0.07 5 5 100000 100 0.00504 0.0096 0.117

57 0.21 5 10 10000 500 0.00252 0.0177 0.135

58 0.21 1 5 10000 500 0.00252 0.0177 0.18

59 0.21 1 10 100000 500 0.00504 0.0096 0.087

60 0.07 1 5 10000 100 0.00252 0.0177 0.16

61 0.21 5 10 100000 100 0.00504 0.0096 0.088

62 0.21 5 10 100000 500 0.00504 0.0177 0.114

63 0.21 5 10 10000 500 0.00504 0.0096 0.131

64 0.07 1 5 100000 100 0.00252 0.0096 0.103
169

Table D-7: EDR factorial design and ANOVA results with labor included. df is the degrees of

freedom. Total degrees of freedom equals 63.

Term Effect Sum of Squares % Contribution F-value Prob > F

A-Electric 0.01191 0.002268 3.00 6650 < 0.0001

B-Brine Disposal 0.01478 0.003496 4.62 10249 < 0.0001

C-Facility Life -0.05103 0.041667 55.1 122156 < 0.0001

D-Flow rate -0.03722 0.022164 29.3 64978 < 0.0001

E-Sodium Removed 0.01128 0.002036 2.69 5970 < 0.0001

F-Membrane Cost 0.00391 0.000244 0.32 716 < 0.0001

G-Labor Cost 0.00784 0.000984 1.30 2886 < 0.0001

AE 0.00228 0.000083 0.11 244 < 0.0001

BE 0.00416 0.000276 0.37 810 < 0.0001

CD 0.01234 0.002438 3.22 7147 < 0.0001

Residuals (df = 53) 0.00031 0.000017 0.02

Total -0.01944 0.075674 100
170

APPENDIX E: UNITS AND UNIT CONVERSIONS

Symbol Definition Equal to:

bbl barrel 42 gal

gal gallon 3.78 L

L liter 0.001 m

3

ft

3

cubic foot 7.42 gal

m

2

square meter 10.76 ft

2

m

3

cubic meter 264.2 gal

gpd gallons per day 1440 gpm

gpm gallons per minute

mgd million gallons per day 1,000,000 gpd

bpd barrels per day 42 gpd

bbl/day barrels per day

kW-hr or kWh kilowatt-hr

$/bbl U.S. $/bbl