ipl engineering design report for cemi427

i

School of Chemical and Minerals Engineering

CEMI427: Conceptual design of uranium extraction plant

G.H. Coetzee 20253362

A. Nel 20311478

D. Postma 20243499

I.S. Scott 20259557

09 November 2009

Engineering Faculty of


ii Declaration

Declaration

We, G.H. Coetzee, A.Nel, D.Postma and I.S. Scott of IPL Engineering hereby declare that

the report entitled:

CEMI427: Conceptual design of uranium extraction plant

submitted for the partial fulfilment of the requirements for the degree B.Eng Chemical

engineering at the North-West University, Potchefstroom campus, is entirely our own work

with external sources referenced in the added reference list.

Sighed at Potchefstroom on the day ______ November 2009.

_______________________________ ________________________________

G.H. Coetzee A.Nel

_______________________________ ________________________________

D. Postma I.S.Scott


iii Executive summary

Executive summary

South Africa is the leading country in regards to nuclear power technology, with regards to

the new nuclear power reactor, PBMR, on the horizon. The new technology pushed the

boundaries of Uranium extraction and enrichment, creating new opportunities for design and

optimization. For the South African market, Eskom is planning to increase their electricity

capacity by implementing only nuclear power (NEA & IAEA, 2007:313).

Uranium is a solid at room temperature and is the last discovered natural occurring element

on the periodic table, with an atomic number of 92. The fact that Uranium is a heavy metal,

makes it a highly toxic element for both human and environmental health. Uranium is a

naturally radioactive substance that emits gamma radiation which is very dangerous, but

very useful if used correctly.

Uranium in the enriched form is mainly used for electricity generation and medical

applications. The medical application is used to treat cancer patients, the gamma rays are

used to break the genetic structure of the cancerous cells. There is industrial use for

Uranium in the X-ray sector. The first use of Uranium was in fact for weapons of mass

destruction, but is considered an ethical aspect in today’s terms.

The world-wide demand for uranium is increasing rapidly and the current production rate is

not able to satisfy this international demand. Although the South African demand is already

satisfied, the export market still holds great advantages. Therefore several uranium mining

and processing plants are being developed or expanded. AngloGold Ashanti is one of the

companies planning to expand its processing of uranium containing ores.

Problem statement

The South Uranium Plant (SUP) currently processes 240 000 ton of ore per month,

producing 624 ton of ADU per year. AngloGold Ashanti is aiming to increase the processed

ore feed to the SUP to 360 000 ton per month, with the additional 120 000 ton of ore per

month. The equipment on the existing plant should be evaluated to determine if the

equipment can handle the increased feed. The existing solvent extraction equipment is

completely exhausted, therefore this unit should be replaced entirely.


iv Executive summary

Process overview The proposed expansion design is able to handle the additional feed and will result in a

production rate of 1068 ton ADU per annum and an overall uranium recovery of 78%. This

design includes a modified process used for leaching which allow the use of the existing

pachuca tanks. Fixed-bed column ion exchange is used rather than CCIX which allows the

use of existing columns and a new solvent extraction (SX) unit is designed to replace current

SX process. The existing precipitation tanks and equipment have enough capacity to handle

the increased feed.

In the leaching process nitric acid is used as oxidant and sulphuric acid is used as lixiviant.

Due to the faster kinetics of nitric acid leaching, only eleven of the existing pachuca tank are

used for the expanded ore feed. Two of the existing unused pachuca tanks are used as

back-up tanks in case of equipment maintenance or breakdown. Each pachuca is air

agitated to ensure sufficient mixing resulting in adequate contact time. The kinetics of nitric

acid leaching is largely dependent on temperature and therefore the first pachuca tank is

heated with steam to above 30 °C with additional steam addition to every fourth tank to

maintain the temperature required. A theoretical analysis on the temperature influence on

the leaching kinetics shows that an increased feed capacity is easily processed by

increasing the leaching temperature or the nitrate addition.

In the counter-current decantation (CCD) process, the leach product is washed to recover

the uranium containing liquids. The existing counter-current decantation equipment are

designed for approximately 8 000 ton ore per day, while an additional 5 000 ton ore per day

capacity is required for the expansion. Therefore an additional counter-current decantation

train is required to process the additional ore feed. Each train consists of six thickeners with

the first thickener in each train used as a clarifier to ensure minimum solids in the solution

sent to ion exchange. The floculant, Magnafloc 90L, is added to the thickeners, except for

the two clarifiers, to promote the sedimentation of the solids. A wash ratio of 1:1 is used to

wash the leach product and results in a uranium recovery of 99.99 %.

The solid slurry product stream from the counter-current decantation is sent to the

neutralization unit where slaked lime is used to increase the pH of the slurry to 10.5 as

required by the gold extraction plant. The neutralization is done in one of the existing

pachuca tanks and after neutralization, the slurry is sent to the gold plant. The liquid product

of the CCD section is sent to ion exchange unit for further processing.


v Executive summary

The pregnant leach liquor, over flow from the clarifiers, proceeds to ion exchange where the

concentration of uranyl sulphate complexes is increased and several of the impurities are

removed. This is achieved by the selective adsorption and elution of the uranyl sulphate

complexes onto resin (Ambersep TM400). The ion exchange unit consists out of five fixed-

bed columns in series with an additional regeneration column. The current adsorption

columns are modified for fixed bed column ion exchange. One important modification is to

enclose the top part of the columns to allow for the back-wash process.

The ion exchange process is divided into four steps; adsorption, washing, elution and

regeneration. The resin remains stationary in the adsorption columns throughout the first

three steps of the process, except for the regeneration step where the resin is moved to the

elution column for treatment. The uranium concentration is increased during the elution step

where diluted sulphuric acid is used to remove uranyl sulphate complexes from the resin.

This results in a concentration of 3.7 g U3O8 per litre in the eluate stream.

The eluate stream is sent to the counter-current solvent extraction section where the

uranium containing liquid is processed to further up concentration and remove impurities.

The solvent extraction process consists of four sections, i.e. extraction, scrubbing, stripping,

and regeneration. The organic solvent feed to the eluate feed has a ratio of 1.1:1. The

extraction section consists of three stages with an additional after settler to reduce solvent

loss. The extractant used is Alamine® 336 with kerosene as diluent and isodecanol as third

phase modifier.

In the scrubbing section the impurities are washed from the solvent with demineralised water

in three mixer-settlers. The demineralised water is recycled to reduce cost with an

occasional purge to remove possibility of build-up in the system. The feed ratio of organic to

demineralised water is 5:1. The stripping section consists of four stages with an additional

after settler to reduce solvent loss. The first mixer-settler is used to reduce the pH of the

stripping solution which results in better pH control. The OK liquor leaving the stripping

section has a concentration of 12 g U3O8 per litre which proceeds to the precipitation section.

The feed ratio of organic to stripping solution is 3.6:1. The organic solvent is continuously

regenerated in the last mixer-settler using a caustic solution with a feed ratio of 1:1.

In the precipitation section the OK liquor and ammonia is fed to the precipitation reactor

where the solid ammonium diuranate (ADU) product is formed. The OK liquor stream is

heated to temperature to 30 °C for adequate reaction kinetics. Two precipitation reactors

are used to acquire efficient precipitation of the product. The density of the slurry product


vi Executive summary

from the second precipitation reactor is increased using one thickener where the over flow is

recycled back to the stripping section of the solvent extraction process. The under flow is

washed in a series of centrifuges to ensure the product specification are met. All the existing

equipment for this section is adequate to handle the increased through-put.

Detail design of solvent extraction unit

In the detail design of the solvent extraction unit the number of stages for each step was

optimised using the McCabe-Thiele method and assumed stage efficiency. The mixers were

designed using basic guidelines for the specific impeller used. The settling kinetics is used

to design dimensions for the settlers. Using the settling kinetics a sensitivity analysis is done

on an estimation of capital cost to obtain an optimum settler vessel dimensions. A

mechanical drawing is done to give guidelines for the construction of the equipment.

Economic evaluation

This expansion project offers an impressive return on investment (ROI) rate of 60.5% with a

payback period of only 3.98 years. The total initial capital investment is calculated as R 847

million which consist of a fixed capital investment of R 720 million and working capital of R

126 million. The operating fixed cost is estimated at R 126 million per annum while the

variable operating cost amounts to R 198 000 per ton ADU produced. The revenue was

determined by assuming R 370 000 will be received for a ton ADU and 50% of the revenue

from the gold sales will be received. The internal rate of return (IRR) is calculated to be

352% with a net positive value (NPV) of R 2.36 billion after a period of 20 years.

Safety consideration

AngloGold Ashanti’s number one values by which the company is driven, is safety and

therefore attention has been dedicated to ensure a safe working environment and minimum

impact on the environment. The general safety design procedure was done by using the

Hazardous and Operability (HAZOP) studies to identify and evaluate the hazards associated

with this uranium extraction plant. The HAZOP studies identified potential risks if accidental

spills of ADU, nitric and sulphuric acid, Magnafloc 90L, kerosene and dust dispersion of the

ore occurs and avoidance measures are suggested. Many of the raw materials are

corrosive and are harmful if humans come into contact with these materials.


vii Executive summary

Special attention was dedicated to the development of fire fighting measures and radiation

exposure control. The solvent extraction unit holds the highest fire risks and therefore

additional fire-fighting equipment is installed. A dead zone of 15 m is created around this unit

to minimize the damage of the other units in case of fire. To eliminate all possible ignition

sources, especially static discharge, all the electrical equipment is sent outside the dead

zone and is covered by zener barriers. Another major safety concern is the exposure to

radiation associated with the handling of uranium ores and products.

The greatest impact of processing plants on the environment is the release of chemical

waste into the atmosphere and water resources. The waste from the SUP include the solid

waste slurry containing the gold, which is neutralized and sent back to the gold plant and the

waste water containing nitrates, which is denitrified by using bio-organisms. A rehabilitation

and decommissioning procedure is included to ensure that the environment is eventually

restored and the impact of the plant on the environment is minimized.

Process control

Control objectives for plant wide control are identified and satisfactory control strategies are

developed with special attention to plant safety. For the solvent extraction unit an

ASPENTech® HYSYS simulation of the proposed control system is done. From this

simulation the feasibility of the control strategies are studied. The proposed control

strategies are sufficient to safely control the process disturbances in order to optimise

production.


viii Table of contents

Table of contents

DECLARATION ii EXECUTIVE SUMMARY iii TABLE OF CONTENTS viii LIST OF FIGURES xiv LIST OF TABLES xvii

CHAPTER 1: INTRODUCTION 1 1.1. Motivation 1 1.2. Problem statement 2 1.3. Assumptions and design limitations 2

CHAPTER 2: LITERATURE STUDY 3

2.1. Ore 3 2.2. Uranium 4 2.2.1. Uses 5

2.2.2. Future tendencies 6

2.2.3. Environmental impact 6

2.2.4. Safety considerations 7

2.3. Ammonium diuranate 8 2.4. Market research 9 2.4.1. Global uranium market 9

2.4.1.1. World demand 10

2.4.1.2. World production 11

2.4.1.3. Relationship between production and demand 11

2.4.2. Southern African market 13

2.4.2.1. South African demand 14

2.4.2.2. South African production 15

2.4.2.3. Relationship between demand and production 16

2.5. Process background 16 2.5.1. Leaching 16

2.5.1.1. Leaching process alternatives 17


ix Table of contents

2.5.1.2. Leaching raw materials 22

2.5.1.3. Leaching design parameters 24

2.5.1.4. Leaching kinetics 25

2.5.2. Ion exchange 27

2.5.2.1. Ion exchange process alternatives 33

2.5.2.2. Ion exchange raw materials 39

2.5.2.3. Ion exchange design parameters 40

2.5.2.4. Ion exchange kinetics and equilibrium 40

2.5.3. Solvent extraction 43

2.5.3.1. Process alternatives 44

2.5.3.2. Raw materials added 45

2.5.3.3. Design parameters 46

2.5.3.4. Kinetics 47

2.5.4. Precipitation 47

2.5.4.1. Precipitation process alternatives 48

2.5.4.2. Precipitation raw materials 50

2.5.4.3. Precipitation design parameters 52

2.5.4.4. Precipitation kinetics 53

2.6. Economic evaluation 55

CHAPTER 3: PROCESS DEVELOPMENT 57

3.1. Level 0: Input information 57 3.1.1. Reactions and reaction conditions 58

3.1.2. Desired production rate and purity 62

3.1.3. Raw materials 62

3.1.4. Processing constraints 64

3.1.5. Other plant and site data 65

3.1.6. Physical properties of all components 65

3.2. Level 1: Batch versus continuous 66 3.2.1. Process units needed 67

3.2.2. Interconnections among units 68

3.2.3. Estimate the optimum processing conditions 69

3.2.4. Additional conceptual design information 70

3.3. Level 2: Input-output structure of the flowsheet 71 3.3.1. Feed purification 71

3.3.2. Recycle streams 72


x Table of contents

3.3.3. Removal and purge streams 72

3.3.4. Number of product streams 72

3.3.5. Preliminary economic potential analysis 74

3.4. Level 3: Recycle structure of the flowsheet 75 3.4.1. Reactor systems required 76

3.4.2. Number of recycle streams 76

3.4.3. Abundance of reactants 78

3.4.4. Operational considerations 78

3.4.5. Recycle economic evaluation 79

3.5. Level 4: Separation systems 80 3.5.1. General structure 80

3.5.2. Vapour recovery system 81

3.5.3. Solid recovery system 81

3.5.4. Liquid recovery system 83

3.5.5. Separation system economic evaluation 85

3.6. Level 5: Heat integration 86 3.7. Equipment design 86 3.8. Mass and energy balance 90

3.8.1. Mass balance 91

3.8.2. Energy balance 91

3.9. Process Flow Diagrams and process description 92 3.9.1. Unit 1 (U01): Leaching, CCD, and neutralization 97

3.9.2. Unit 2 (U02): Ion exchange 99

3.9.3. Unit 3 (U03): Solvent extraction 101

3.9.4. Unit 4 (U04): Precipitation 103

3.10. Innovations 105

CHAPTER 4: DETAIL DESIGN 108 4.1. Choice of system type 108 4.2. Kinetics and thermodynamics 110 4.3. Detail chemical design 112

4.3.1 Number of stages 112

4.3.2 Mixer 113

4.3.3 Settler 114

4.3.3 Pipe sizing 115

4.3.4. Pump sizing 116


xi Table of contents

4.3.5. Control valve sizing 117

4.4. Start-up and shut-down procedures 119 4.4.1. Commissioning of process unit 119

4.4.2. Shutdown procedures 121

4.4.3. Emergency shutdown procedures 121

4.5. Mechanical aspects 123

CHAPTER 5: TECHNO-ECONOMIC EVALUATION 130

5.1. Definitions and assumptions 130 5.1.1. Definitions 131

5.1.2. Assumptions 134

5.2. Estimation of capital, operating cost and revenue 134 5.2.1. Capital investment 135

5.2.2. Operating cost 137

5.2.3. Revenue 138

5.3. Cash flow analysis 138 5.4. Economic sensitivity analysis 140 5.5. Recommendation for profitability 142

CHAPTER 6: SAFETY AND ENVIRONMENT 143

6.1. Overall safety specifications 143 6.1.1. Leaching 144

6.1.2. Ion exchange 145

6.1.3. Solvent extraction 145

6.1.4. Precipitation 146

6.1.5. Fire fighting 146

6.1.6. Training and personal safety 148

6.1.7. Danger zones and signs 149

6.1.8. Emergency response plan 150

6.1.9 Radiation control 151

6.2. HAZOP level 1 152 6.2.1. Project definition 153

6.2.2. Process description 154

6.2.3. Assessment of chemical hazards 155

6.2.4. Assessment of chemical interactions 156


xii Table of contents

6.2.5. Means of handling 158

6.3. Environmental impact and management 159 6.3.1. Waste block diagram analysis 159

6.3.2. Handling and disposal of wastes 162

6.3.3. Accidental releases of hazardous materials 164

6.3.4. Rehabilitation and decommissioning of plant 164

6.4. Plant layout and positioning 165

CHAPTER 7: PROCESS CONTROL 167

7.1. Plant wide control 167 7.1.1. Unit 1: Leaching, CCD and neutralization 168

7.1.2. Unit 2: Ion exchange 175

7.1.3. Unit 3: Solvent extraction 179

7.1.4. Unit 4: Precipitation 182

7.2. Specific safety considerations for solvent extraction 188 7.3. Detailed process control for solvent extraction 199

7.3.1. Flow control loops 201

7.3.2. Tank level control loops 203

7.3.3. pH control loops 205

7.3.4. Range, alarms and trips 207

7.4. Dynamic control analysis 209 7.4.1. Tuning control loops 210

7.4.2. Variable pairing 216

7.5. Conclusion 218 REFERENCES 219

APPENDIX A: MASS AND ENERGY BALANCE 229 APPENDIX B: EQUIPMENT SIZING 254 APPENDIX C: DETAIL EQUIPMENT CALCULATIONS 264 APPENDIX D: TECHNO-ECONOMIC CALCULATIONS 283 APPENDIX E: PLANT LAYOUT AND POSITIONING 299 APPENDIX F: MSDS INFORMATION 302


xiii List of figures

List of figures

CHAPTER 2: LITERATURE STUDY Figure 2.1: Processed yellow cake 8

Figure 2.2: Global energy consumption 9

Figure 2.3: Distribution of global uranium requirements for 2006 11

Figure 2.4: Cumulative global uranium requirements and demand 13

Figure 2.5: Gold price for 2005 to 2009 14

Figure 2.6: The classification of the various leaching techniques 18 Figure 2.7: In-situ leaching process description 19

Figure 2.8: Illustration of a heap leach process 20 Figure 2.9: Typical adsorption 30

Figure 2.10: Eluate concentration profile 32

Figure 2.11 Fixed bed ion exchange column layout 35

Figure 2.12: Typical solvent extraction process 44

CHAPTER 3: PROCESS DEVELOPMENT

Figure 3.1: Simplified block flow diagram of process 68

Figure 3.2: Input-output structure of overall process 73

Figure 3.3: Block flow diagram for reactor system 76

Figure 3.4: Recycle structure 77

Figure 3.5: General separation structure 81

Figure 3.6: Waste solid recovery system 82

Figure 3.7: Solid recovery system for ADU 83

Figure 3.8: Schematic representation of ion exchange 84

Figure 3.9: Schematic representation of solvent extraction 84

CHAPTER 4: DETAIL DESIGN

Figure 4.1: Representation of mixer-settler equipment 109

Figure 4.2: Stripping loading isotherm with (NH4)2SO4 111

Figure 4.3: Back-up pump configuration 117

Figure 4.4: Pumper mixer Impeller 124


xiv List of figures

Figure 4.5: pH box for measurement 125

CHAPTER 5: TECHNO-ECONOMIC EVALUATION

Figure 5.1: Cash flow diagram 139

Figure 5.2: Economic sensitivity analysis 141

CHAPTER 6: SAFETY AND ENVIRONMENT Figure 6.1: Examples of necessary signage 150

Figure 6.2: HAZOP structure for plant 153

Figure 6.3: General waste block diagram structure 159

Figure 6.4: Leaching waste block diagram 160

Figure 6.5: Ion exchange waste block diagram 160

Figure 6.6: Solvent extraction waste block diagram 161

Figure 6.7: Precipitation waste block diagram 161

Figure 6.8: Plant layout 166

CHAPTER 7: PROCESS CONTROL

Figure 7.1: Control schematic of the leaching section 169

Figure 7.2: Control schematic of the counter current decantation section 172

Figure 7.3: Control schematic of the neutralization process 174

Figure 7.4: Control schematic of the adsorption stage of ion exchange 176

Figure 7.5: Control schematic of the back-wash stage of ion exchange 177

Figure 7.6: Control schematic of the elution stage of ion exchange 178

Figure 7.7: Control schematic of the regeneration stage of ion exchange 179

Figure 7.8: Basic control schematic of the solvent extraction unit 180

Figure 7.9: Control schematic of the precipitation reactors in the precipitation unit 183

Figure 7.10: Control schematic of the solid-liquid separation in the precipitation unit 186

Figure 7.11: Extraction simulation flowsheet 211

Figure 7.12: Ratio control results in extraction section 212

Figure 7.13: Scrubbing simulation flowsheet 213

Figure 7.14: The primary control loop results 214

Figure 7.15: The secondary control loop results 215


xv List of figures

APPENDIX A: MASS AND ENERGY BALANCES

Figure A.1: Aspen Tech® simulation 230

APPENDIX B: EQUIPMENT SIZING Figure B.1: Influence of nitrate concentration on reaction kinetics 257

Figure B.2: Influence of temperature on reaction kinetics 258

Figure B.3: Leach tank size evaluation 258

APPENDIX C: DETAIL DESIGN CALCULATIONS

Figure C.1: Modelling of the extraction loading isotherm 264 Figure C.2: Loading isotherm data for the stripping section 265

Figure C.3: Solvent extraction process flow schematic 265

Figure C.4: McCabe-Thiele for the Extraction section 267

Figure C.5: McCabe-Thiele for the Stripping section 268

Figure C.6: Vessel width sensitivity analysis 273

Figure C.7: Dispersion layer velocity sensitivity analysis 274

Figure C.8: Economical sensitivity analysis on the extraction settlers 275

Figure C.9: Dispersion height profile over the length of the extraction vessel 276

Figure C.10: Economical sensitivity analysis on the stripping settlers 278

Figure C.11: Height profiles for internal recycle and without internal recycle 279

APPENDIX D: TECHNO-ECONOMIC CALCULATIONS

Figure D.1: Cost of general-purpose centrifugal pumps 285

Figure D.2: Cost of electric motors 285

Figure D.3: Installed cost for single-compartment thickeners 286

APPENDIX E: PLANT LAYOUT AND POSITIONING

Figure E.1: Google Earth air photo 299

Figure E.2: Existing plant layout of South Uranium Plant with number legend 300

Figure E.3: Existing plant layout of South Uranium Plant with colour legend 301


xvi List of tables

List of tables

CHAPTER 2: LITERATURE STUDY

Table 2.1: Physical and chemical properties of uranium 4

Table 2.2: Proposed interim water quality guidelines for the radiation dose in

drinking water 7

Table 2.3: Uranium plants in South Africa 15

Table 2.4: Typical composition of Rand leach liquor 27

Table 2.5: Suggested concentrations for eluant choices 31

Table 2.6: Feed and product streams for ion exchange 39

Table 2.7: Comparison of solvent extractor types 45

Table 2.8: Comparison of neutralization reagents 51

Table 2.9: Summary of uranium project survey 56

CHAPTER 3: PROCESS DEVELOPMENT Table 3.1: Leaching reactions 58

Table 3.2: Reaction information for ion exchange 59

Table 3.3: Reaction information for solvent extraction 60

Table 3.4: Reactions kinetics and conversion 61

Table 3.5: Ore composition 63

Table 3.6: Summary of raw materials 64

Table 3.7: Utilities 65

Table 3.8: Physical properties of unknown compounds 66

Table 3.9: Processing conditions 69

Table 3.10: Product stream classification 74

Table 3.11: Level 2 preliminary mass balance with prices 75

Table 3.12: Capital and operation costs for reactors 79

Table 3.13: Separation systems cost calculations 85

Table 3.14.a: Summary of leaching equipment 87

Table 3.14.b: Summary of counter-current decantation equipment 88

Table 3.14.c: Summary of ion exchange equipment 89

Table 3.14.d: Summary of precipitation equipment 90

Table 3.15: Overall mass balance 91


xvii List of tables

Table 3.16: Innovations 106

CHAPTER 4: DETAIL DESIGN

Table 4.1: Extraction loading isotherm constants 113

Table 4.2: Results from McCabe-Thiele method 113

Table 4.3: Mixer box design results 114

Table 4.4: Designed dimensions for the settler vessels 115

Table 4.5: Solvent extraction pipe sizes 116

Table 4.6: Pump sizing results 116

Table 4.7: The calculated pressure drop and valve coefficients 118

CHAPTER 5: TECHNO-ECONOMIC EVALUATION

Table 5.1: Multiplying factors used in the delivered-equipment cost method 136

Table 5.2: Operating costs 137

Table 5.3: Revenue from product and by-product sales 138

Table 5.4: Profitability results from cash flow analysis 139

CHAPTER 6: SAFETY AND ENVIRONMENT

Table 6.1: Examples of radiation dosages 152

Table 6.2: Projected production rate for the upgraded plant 153

Table 6.3: List of chemicals 155

Table 6.4: Chemical hazard data sheet (HS1A) 156 Table 6.5: Chemical interactions data sheet (HS1B) 157

Table 6.6: Means of handling (HS1C) 158

Table 6.7: Waste classification 162

Table 6.8: Accidental releases 164

CHAPTER 7: PROCESS CONTROL Table 7.1: HAZOP 3 Proforma HS3A form 191

Table 7.2: HAZOP 3 record for Proforma HS3A 192

Table 7.3: Specifications for flow controllers 203

Table 7.4: Specifications for drain system level control 204


xviii List of tables

Table 7.5: Specifications for feed flow level control 205

Table 7.6: Specifications for pH control 206

Table 7.7: RAT list for the solvent extraction section 207

Table 7.8: Control loop tuning values 216

Table 7.9: Relative gain array for control on U03-ST07 217

APPENDIX A: MASS AND ENERGY BALANCES Table A.1: Summary of mass balances for individual units 229

Table A.2: Mass balance for leaching process 231

Table A.3: Mass balance for CCD process 234

Table A.4: Mass balance for neutralization process 237

Table A.5: Mass balance for ion exchange process 240

Table A.6: Mass balance for solvent extraction process 243

Table A.7: Mass balance for precipitation process 246

Table A.8: Leaching pachucas energy balance 249

Table A.9: Counter-current decantation energy balance 251

Table A.10: Neutralisation energy balance 252

Table A.11: Precipitation energy balance 253

APPENDIX B: EQUIPMENT SIZING

Table B.1: Leaching kinetics constants 254

APPENDIX C: DETAIL DESIGN CALCULATIONS

Table C.1: Variables for Equation C-22 and C-23 282

APPENDIX D: TECHNO-ECONOMIC CALCULATIONS Table D.1: Purchased equipment cost 287

Table D.2: Capital investment calculations 288

Table D.3: Detailed operating cost calculation 290

Table D.4: Revenue calculations 292

Table D.5: Cash flow analysis 292

Table D.6: Chance in variables 298

Table D.7: Economic sensitivity analysis in terms of percentage rise/fall of NPV 298


1 Chapter 1: Introduction

Chapter 1: Introduction

Uranium is a remarkable element that can be used as an abundant source of concentrated

nuclear energy. Although the radioactivity and possible application in weaponry made the

world sceptical about the uses of uranium, the recent developments in nuclear power

stations and reactors have created an entire new market for uranium.

Uranium was discovered in 1789 by Martin Heinrich Klaproth just after the discovery of the

planet Uranus, from which the name for uranium derives. It appears that uranium was

formed as a decay product of elements with a higher atomic weight, in supernovae about 6.6

billion years ago. Uranium is widely available in most rocks on earth and the radioactive

decay of uranium is a main heat source inside the earth (Cleveland, 2008).

In nature, uranium is mainly found as an oxide such as Triuranium octaoxide (U3O8) and

uranium dioxide (UO2), but can also occur in the form of uranium hexafluoride (UF6) and

uranium tetrafluoride (UF4). Uranium is also found in the form of a metal which is more

dense, more ductile and softer than steel, with a density of 19 000 kg/m3 (Cleveland, 2008).

The first use of uranium dates back to 79 A.D., when it was used to produce yellow-coloured

glass. Presently uranium is used mainly to generate energy, in research reactors and can

also be used to create explosives and weaponry. Uranium is essential for the nuclear

enterprise and the rapid development of nuclear reactors and power plants results in an

increasing demand for uranium (Cleveland, 2008).

1.1. Motivation

Uranium is a by-product of gold extraction which ends up in the mine tailings which has an

enormous environmental impact. The pre-leaching of the gold ore removes the uranium and

some of the impurities while increasing the gold recovery. Further more this results in a

great economic potential due to the increased gold recovery and uranium by-product. The

global demand for alternative energy sources has placed strain on the global uranium supply

and stockpiles. Due to the economic and environmental advantages associated with

uranium extraction, it is necessary to conduct a feasibility study on the expansion of the

South Uranium Plant (SUP) at AngloGold Ashanti near Orkney.

http://www.eoearth.org/article/Klaproth%2C_Martin_Heinrich�


2 Chapter 1: Introduction

1.2. Problem statement

The existing South Uranium Plant (SUP) should be upgraded to handle an additional feed of

120 000 ton ore per month which increases the total feed to 360 000 ton ore per month. The

additional ore feed is received from another gold plant and will have slightly different

specifications as the regular feed. The preliminary economic evaluation showed that the

following is needed:

• Upgrade of the leaching section.

• Upgraded or new ion exchange section.

• New solvent extraction section.

• Upgrade of precipitation section.

Information given:

• The upgraded equipment will be situated on the existing SUP.

• The final product i.e. yellow cake (ADU) at a mass percentage of 35% U3O8 that

should be stored for further shipping.

• All utilities and certain raw materials are provided at the existing plant at a fixed cost.

• Gas and liquid emissions should adhere to applicable laws and regulations and

should be processed on site, but process equipment does not have to be designed.

1.3. Assumptions and design limitations

All the assumptions made throughout the feasibility study is given in the respective sections.

The assumptions are as far as possible based on literature sources, existing plant data, and

logical reasoning. The aspect falling outside the battery limits for the design project are as

follows:

• Crushing and milling of ore received from gold plant.

• Gold processing.

• The delivery of utilities and raw materials.

• The treatment of wastes.

• Transport of ADU product to NUFCOR.


3 Chapter 2: Literature survey

Chapter 2: Literature survey

This chapter gives a brief oversight into the processing of uranium and describes

subsequent effects on the environment, population and economics. In this chapter the

following aspects are addressed:

• Ore that is used in the extraction of uranium.

• Uranium, its properties and uses.

• The product ammonium diuranate.

• The market research for uranium.

• Background on uranium processing processes.

• Economic evaluation for uranium extraction from literature.

Understanding and respecting the above aspects will lead to sustainable development,

which is the most important aspect to consider.

2.1. Ore

The ore used at the AngloGold Ashanti Southern Uranium Plant is mined from the

Witwatersrand basin. The ore is primarily used for gold extraction, with uranium as a main

co-product. The ore from the Witwatersrand Basin in South Africa is a quartz-pebble deposit

type (British Geological Survey, 2007: 5).

Quartz-pebble deposit type is formed from ancient sedimentary deposits buried when the

atmosphere was believed to be less oxidizing than today. From literature it is also believed

that the rapid filling of the basin by rivers isolated the uranium before oxidation could take

place. The typical grade of the uranium in the quartz-pebble deposit type ore is 130 to 1100

ppm uranium (British Geological Survey, 2007: 7).

The ore sent to AngloGold Ashanti South Uranium Plant, consist of the following non

uranium containing minerals (Lottering et al., 2007: 18)



• Quartz (SiO2).

• Muscovite (K-Al-silicate).

• Chlorite (Mg-Fe-Al-silicate).

• Pyrophylite (Al4(Si8O2)(OH)4).

• Pyrite (FeS2).

• Albite (Na-Al-silicate).

As well as the above minerals, the uranium containing minerals of the ore are (Lottering et

al., 2007: 18):

• 84.9% Uraninite (UO2).

• 12.9% brannerite ((U,Th,Ca)(Ti,Fe)2O6).

• 2% U-phosphate ((U,Cl)PO4).

• 0.2% coffinite (U(SiO)41-x(OH)4x).

Uraninite is the mineral from which it is the simplest to extract uranium from, and the

extraction chemistry will be based on this mineral.

2.2. Uranium

Uranium is the last natural occurring metal in the periodic table and is ranked 49th most

abundant element in the earth’s crust out of 92 elements. Uranium was discovered in 1789

by Martin Heinrich Klaproth in Berlin, German, while investigating the mineral pitchblende.

Uranium is a gray metal which is chemically reactive. Table 2.1 gives some of uranium’s

physical and chemical properties (Enghag, 2004: 1166-1169).

Table 2.1: Physical and chemical properties of uranium

Property Information

Symbol U

Molar weight (g/mol) 238.03

Density (kg/m3) 18 950

Melting point (K) 1405.5

Isotope range, all nuclides radioactive 218 to 242



The history and properties of uranium is important, but will serve no purpose if there is no

use for the element. In the next sub-section the uses of uranium is discussed.

2.2.1. Uses

Uranium has various uses, with its main use in the generation of electricity. It is estimated

that 95% of all uranium is used to generate electricity, with the remaining 5% aimed at other

uses (British Geological Survey, 2007: 15).

To create electricity from uranium, nuclear power stations use enriched uranium-235 as heat

source to convert water into steam. The steam produced from this is used to turn turbines

which generate the electricity. This principle is the same as fossil fuel power stations, but

the fossil fuel is replace by a significantly lower amount of uranium and with little to none

emissions (British Geological Survey, 2007: 15).

Some of the other uses of uranium include the following (British Geological Survey, 2007:

18-19):

• Nuclear-powered ships.

• Research reactors.

• Desalination.

• Weapons.

The South African Nuclear Energy Corporation (NECSA) has a research nuclear reactor

situated in Pelendaba, named the Safari-1 reactor. Currently this reactor is the leading

producer of the medical isotope molybdenum-99 in the world. One of the other main uses of

the reactor is the irradiation of silicon semiconductors (NECSA, 2009).

For uranium to be used in all these practices, there needs to be a supply in the future, the

next sub-section will discuss the future tendencies of uranium production in South Africa.



2.2.2. Future tendencies

With the rise of concern about the global environment and global warming, uranium is set to

replace fossil fuels in the generation of electricity. With the increase in nuclear reactors, the

consumption of uranium will increase, which must lead to an increase in production of

uranium.

In 2009 the production of U3O8 in South Africa is proposed to be 2 800 tons per annum.

Furthermore it is estimated that the production of U3O8 will surpass 5 000 tons per annum in

2011. This value will be reached if all the proposed projects start on schedule with the

proposed production rates (Damarupurshad, 2007: 11).

2.2.3. Environmental impact

Uranium is naturally spread widely throughout the environment and can be found in small

amounts in rock, soil, air and water. Humans release additional uranium metals and

compounds into the environment as a result of milling and mining which causes

environmental and health concerns (Lenntech, 2008). Uranium poses two threads that will

be investigated, its toxicity and the effects of radioactivity.

Uranium itself is radioactive, though with the major isotope U-238 having a half-life equal to

the age of the earth; it is certainly not strongly radioactive (World Nuclear Association, 2008).

Human activities such as mining enhance the transport of radionuclides into the environment

by transferring the ore from underground to tailing dams on the surface. These

radionuclides can separate from the bulk ore and pass into the environment via water

streams, for example, uranium in underground water pumped to the surface, and uranium

leakage from tailing dams resulting from oxidation of pyrite (Wendel, 1998:92).

Uranium is a heavy metal that is also toxic chemically. The toxicity becomes a danger when

it is found in groundwater contaminating soil where crops are grown, and in drinking water.

As water, containing oxygen, passes over rocks and soil, many compounds and minerals,

such as uranium, dissolve and is transferred into the groundwater. Uranium in groundwater

also results from human activities such as mining, combustion of coals and other fuels, the

use of phosphate fertilizers, and nuclear power production (Skipton et al 2008:1). The table

below shows the interim water quality guidelines for the radiation dose in drinking water

proposed by the Department of Water Affairs & Forestry (Kempster, 1999).



Table 2.2: Proposed interim water quality guidelines for the radiation dose in drinking water

Radiation dose (nSv/a) Suitability

<= 0.1 (WHO reference level) Ideal, suitable for lifetime use

> 0.1 and <= 0.25 Water acceptable for lifetime use, subject to

confirmation of dose.

> 0.25 and <= 1 Water acceptable for short term use. Use in

longer term (lifetime) requires further

investigation.

> 1 and <= 5 Unacceptable for lifetime use

> 5 Unacceptable even for short term use

The health risks associated with uranium include kidney diseases, diseases of the

respiratory tract and cancer. Uranium also binds to biological molecules, is distributed within

the body and builds up in bone and teeth (Lindemann, 2008:5).

2.2.4. Safety considerations

To provide a safe working environment it is important to identify the potential health hazards,

provide the correct first aid measures and accidental release measures, have appropriate

handling and storage procedures and have preventative exposure controls and personal

protection.

The potential hazards identified when working with uranium include acute and chronic health

effects. It can be corrosive and irritant when in it comes in contact with the skin or eyes.

Liquid or spray mist which is ingested or inhaled may produce tissue damage particularly on

mucous membranes of eyes, mouth and respiratory tract. Inhalation may also produce

severe irritation of respiratory tract, characterized by coughing, choking or shortness of

breath. Repeated or prolonged exposure can produce accumulation and damage in target

organs. Severe over-exposure can result in death (Science Lab. 2008).



There must be first aid measures for eye contact, skin contact, ingestion and inhalation both

minor and serious. If there is a small spill it must be diluted with water, absorb with an inert

dry material and place in an appropriate waste disposal container. If there is a larger spill

the MSDS must be consulted for correct clean-up method. The uranium must be kept in a

dry container and should be stored in a separate safety storage cabinet or room. When

working with uranium personal protection include gloves, boots, face shield, full suit and

approved/certified vapor respirator (Science Lab, 2008).

For additional safety measurements, employees can be monitored to ensure the uranium

levels in their bodies are under the acceptable limit.

2.3. Ammonium diuranate

Yellow cake is the universal name for ammonium diuranate (ADU), which is produced at

most uranium extraction plants. ADU’s chemical formula is (NH4)2U2O7, which is yellow in

color and has a melting point of 2878 °C (Hausen, 2009). Figure 2.1 is a picture of

processed yellow cake (Snooper, 2008).

Figure 2.1: Processed yellow cake

In South Africa, the yellow cake produced is sent to NUFCOR for processing consisting of

filtering, drying and calcining. After processing the yellow cake is shipped to international

countries for use in nuclear reactors and for its other uses.



2.4. Market research

As with most commodities, the price of uranium is greatly influenced by the supply and

demand. It is therefore very important to determine the current state of the uranium market

which will shed light on the present and future economic viability of uranium production.

Since the extraction of uranium from the gold ore holds recovery benefits for the gold

extraction, the gold market outlook also has an influence on the economic viability of

uranium production. The South African and global markets are investigated and compared

in the following section.

2.4.1. Global uranium market

The climate changes, that are observed globally, cause increasing concern for the release of

greenhouse gasses such as carbon dioxide and methane. For this reason legislation is

implemented to decrease the use of fossil fuels for energy. As mentioned before, uranium is

mainly used in the nuclear energy industry and this is a very attractive alternative fuel

compared to fossil fuels. This results in an increasing global demand for uranium.

Figure 2.2 shows the world energy use from 1980 to 2006 and what type of source the

energy came from as derived from tables produced by the Energy Information Administration

(EIA) (2008).

0

20

40

60

80

100

120

140

Gig

awat

t ho

urs

Mill

ions

Petroleum Dry Natural Gas Coal Net Hydroelectric Power Net Nuclear Electric Power Others

Figure 2.2: Global energy consumption



As seen in Figure 2.2, the world has become increasingly dependent on nuclear energy as a

source of energy. The world’s energy demand is increasing with a steady fraction of energy

obtained from nuclear power plants.

2.4.1.1. World demand

Since the world nuclear energy consumption is increasing, the global uranium demand is

also increasing. There was however a drop in world uranium usage from 66 500 tU in 2007

to 64 500 tU in 2008. This drop in uranium usage was due to the closure of seven units in

Japan’s largest nuclear power plant (Kashiwazaki-Kariwa) and other nuclear power plants

were closed for maintenance and repairs. It is reported by the Department: Minerals and

Energy of South Africa (2008:52) that in 2007, uranium was used as fuel in 439 nuclear

power plants around the world and that there are 36 currently under construction. This

means that the uranium demand will increase when these new nuclear power plants are

finished (Department: Minerals and Energy of South Africa, 2008:52).

There are other factors that may decrease the overall global uranium demand. The

decommissioning of old or first generation nuclear power plants will eventually start causing

a decrease in uranium demand. The efficiency of nuclear power plants is also constantly

improved through new research and technologies, which decreases the uranium demand.

Regardless of these small factors, the global uranium demand is expected to increase since

the world electricity requirements are increasing rapidly (World Nuclear Association, 2008).

Since the planning and construction of nuclear power plants is expensive, developed

countries are the largest uranium consumers. However, nuclear power is a very attractive

source of energy for developing countries with a lack in infrastructure, such as African

countries, since small amounts of fuels is necessary and transport is easier. Figure 2.3

shows the geographical distribution of global uranium requirements in 2006 (NEA & IAEA,

2007:52).



Figure 2.3: Distribution of global uranium requirements for 2006

As seen from Figure 2.3, the developed world uses the most uranium. The United States of

America is the largest user of uranium followed by France, Japan and then Germany

(Department: Minerals and Energy of South Africa, 2008:53). The United States and

Canada are large producers of uranium and therefore provide for their own demand.

France has many nuclear power plants but virtually no natural uranium resources. Most of

France’s 10 500 tU needed per year for electricity generation is imported from Canada and

Niger. France do supply a part of its own uranium demand through conversion and

enrichment plants that recycle used uranium. This causes 30% more energy produced from

the original uranium and a great reduction in the amount of nuclear waste disposed (World

Nuclear Association, 2009).



2.4.1.2. World production

In 2008 the world produced 44 000 tU which is quite higher than the 41 300 tU in 2007. The

largest uranium producers are Canada, Kazakhstan and Australia which jointly produced

approximately 60% of the global uranium production in 2008. The forecast global uranium

production for 2009 is 49 400 tU according to the World Nuclear Association (2009). This

sharp increase is expected because of major increases in production of the leading uranium

producing countries.

The methods of uranium production have also changed over the last few years. In situ

leaching makes the extraction of uranium much more economically viable because the ore

does not have to be mined, crushed and milled. This makes the In situ leaching method an

increasingly attractive choice for the production of uranium. Currently 28% of the uranium

produced is produced with the In situ leaching method (World Nuclear Association, 2009).

2.4.1.3. Relationship between production and demand

From the figures in the section above it is seen that in 2008 the uranium production rate only

provided for approximately 70% of the global demand. The rest of the uranium demand is

currently supplied for with uranium stockpiles. These stockpiles were built up during the cold

war when the United States and Russia produced high-enriched uranium (HIU) for the

manufacturing of nuclear weapons. These weapons are now being dismantled and the

uranium used in commercial nuclear reactors. In 2006 the uranium from the military

provided for more than 50% of the nuclear fuel in United States nuclear reactors (Johnson,

2007).

There is concern for when these stockpiles are depleted, because the current production

rate is not enough for the annual uranium demand. This is seen in Figure 2.4 which shows

the cumulative global uranium requirements and demand (NEA & IAEA, 2007:75).



Figure 2.4: Cumulative global uranium requirements and demand

From the slopes of the end of the graph in Figure 2.4, it is derived that at some point in the

future the global demand will exceed the global supply of uranium. If the global uranium

production increases as it is currently doing, then the uranium supply will be able to sustain

the growing demand. Another obstacle in uranium production is the availability of

enrichment facilities. There are strict regulations for the creation of enrichment facilities to

ensure it is only used to produce low-enriched uranium (LEU) and not HEU which is readily

used to construct nuclear weapons.

2.4.2. Southern African market

Only South Africa and Namibia are currently producing uranium in Southern Africa. The

Rossing mine in Namibia is the world’s third largest uranium producing mine and produced 3

400 tU in 2008. The uranium demand of these countries is very low which means that most

of the produced uranium is exported. In South Africa uranium is produced as a by-product of

the gold extraction process and causes an improved gold recovery. Therefore the uranium

extraction from gold ore has an increased economical potential which is greatly dependant

on the gold price.

In 2008 the world economy entered a crisis period which caused many commodity prices,

including that of gold, to fall. In Figure 2.5 the gold price over the last five years is shown

(Goldprice™, 2009).



Figure 2.5: Gold price for 2005 to 2009

The effect of this economy crisis on the gold price is clearly visible in Figure 2.5. However

Figure 2.5 also shows that the gold price did recover and is stable. This provides

economical potential and security for the production of uranium as a by-product of gold

extraction.

2.4.2.1. South African demand

In Southern Africa there are only three nuclear reactors with a very low nuclear fuel

requirement. Two are situated at the Koeberg nuclear power plant near Cape Town, South

Africa. These two reactors consume only 292 tU per year and provide 5% of South Africa’s

electricity (NEA & IAEA, 2007:313). Since Southern Africa does not have any operating

uranium enrichment facility, all of its enriched uranium fuel has to be import (World Nuclear

Association, 2009).



Eskom is South Africa’s leading electricity company, providing 95% of its electricity and 60%

of Africa’s electricity. In 2008 South Africa’s electricity demand exceeded the supply and

wide spread power failures became a great concern. This electricity crisis prompted Eskom

to drastically increase its capacity. Eskom’s initial plans to increase its nuclear capacity to

20 GWe by 2025 have been halted due to a lack of finances. New projections by Eskom are

that their nuclear capacity might reach 6 GWe by 2025 through the construction of three new

nuclear power plants (NEA & IAEA, 2007:313).

2.4.2.2. South African production

In 2007 South Africa produced 525 tU in 2007 of which Anlogold Ashanti Ltd. produced the

greatest part. This company was the only uranium producer in South Africa for a few years.

However, the recent increase in the uranium price has made the production thereof more

economically viable and more uranium producing plants are being constructed and re-

opened in South Africa. Table 2.3 shows a list of the uranium plants in South Africa that is

currently operating or will be operating in the near future as derived from the World Nuclear

Association (2009).

Table 2.3: Uranium plants in South Africa

Plant/Project Company Projected capacity (tU/yr) Reached in

South Uranium Plant AlgloGold Ashanti Ltd. 760 2012

Dominion Reefs Uranium One 1470 2011

Ryst Kuil UraMin Inc. 1150 2009

Ezulwini First Uranium Corp. 116 2010

Buffelsfontein First Uranium Corp. 330 2010

It is seen from Table 2.3 that the uranium production in South Africa will most probably

increase over the next few years with many expansions and new plants. The Nuclear Fuel

Corporation of South Africa (Nufcor) does the final processing step in the uranium production

process and has a capacity of 3 400 tU per year (NEA & IAEA, 2007:313). From this it is

seen that if all the planned plant do reach full production capacity, Nufcor will not be able to

handle all the uranium at its present capacity. However some of the above mentioned

companies were very optimistic in their projections.



2.4.2.3. Relationship between demand and production

As mentioned before South Africa currently has very low uranium requirements which may

increase in future as new nuclear power plant is constructed. However South Africa does

not have any enrichment facilities, which means the nuclear reactor fuel is imported. The

Nuclear Energy Corporation of South Africa (Necsa) did enrich uranium for research, military

use and the Koeberg Nuclear power plant, but has seized these operations according to the

agreement under the Non-Proliferation Treaty. Therefore all uranium produced in South

Africa is exported to be enriched and all nuclear reactor fuel is either imported or supplied by

stockpiles from the military (World Nuclear Association, 2009).

2.5. Process background

The process used for the extraction of uranium is described in detail in the form of a

literature study. The processes described are leaching, solid-liquid separation, ion-

exchange, solvent extraction and precipitation. Each step in the process is described in the

following structure:

• A brief description of the background theory.

• The process alternatives and the safety and environmental aspects.

• The raw materials added to the process.

• The design parameters.

• The kinetics present in the process.

The processes will be evaluated further while new concepts are developed and important

new information can lead to the useless rendering of the process.

2.5.1. Leaching

Leaching is the first hydrometallurgical process in the extraction of uranium, hydrometallurgy

is the process where metals are produced from an aqueous solution (Minerals counsil of

Australia, 2006: 5). Leaching is where the valuable minerals are extracted from grounded

ore by dissolving, preferably selective, the mineral in an aqueous solution. The chemical

used to dissolve the minerals is known as the leaching agent (Lunt & Holden, 2006: 4).



From the above definitions, it is easy to understand the importance and economical impact

of leaching. Since leaching is the first extraction stage in the processing of minerals, it is

important to selectively dissolve most of the essential minerals. The valuable minerals that

are not dissolved in the leaching section are lost in the tailings. These minerals will not only

be lost as an economical resource but it can also have devastating ecological effects.

It is possible to economically extract more than one valuable metal from the ore, in South

Africa this is done with ore that contains both uranium and gold. In the case of gold and

uranium the different minerals are removed using separate leaching systems. Leaching is

the most expensive key component in the capital cost of an uranium processing plant, the

capital cost of acid leaching equipment can contain up to 49% of the capital cost expenses

of the project (Lunt & Holden, 2006: 4).

2.5.1.1. Leaching process alternatives

The driving force for the development of different leaching techniques is optimization of the

leaching process. The five main factors to take into account when choosing a leaching

process are (Minerals counsil of Australia, 2006: 170):

• The degree of recovery for the desired metal.

• The selectivity required for the desired metal.

• Leaching time required to achieve the desired recovery.

• The capital cost required for the leaching equipment.

• The operating cost of the reactants.

The optimization of the leaching process has lead to numerous techniques. The leaching

operation can be done in batch, counter current, atmospheric or above atmospheric

pressure and ambient or elevated temperatures. It is therefore important to classify the

techniques and this is illustrated in Figure 2.6 (Gupta,2003: 480):



Figure 2.6: The classification of the various leaching techniques

Leaching techniques are separated into two main groups, percolation leaching and agitation

leaching. In percolation leaching, the ore is static and the leaching agent is percolated

through the ore to allow dissolution. In agitation leaching, the ore is suspended in the

leaching agent to allow dissolution. The dissolution kinetics for percolation leaching is very

slow but it requires less capital cost when compared to agitation leaching (Gupta,2003: 480).

Percolation leaching is discussed first.

In-situ leaching

The first percolation leaching technique discussed is In-situ leaching. The main use of In-

situ leaching is applied in the extraction of uranium. In-situ leaching has a minimal surface

disturbance and is therefore an environmentally friendly mining process. In-situ leaching

reverses the natural process which deposits the uranium in the porous sandstone (Minerals

counsil of Australia, 2006: 170). Figure 2.7 illustrates the arrangement of an in-situ leaching

operation (NunnGlow, 2008).



Figure 2.7: In-situ leaching process description

The confined deep aquifer from Figure 2.7 contains the mineral that is leached, the aquifer

should be permeable to the leaching agent. The aquifer is between two clay layers, the clay

layers restrain the leaching agent inside the aquifer and thus preventing the leaching agent

from contaminating the water table. The leaching agent is injected to the aquifer through the

injection wells and the pregnant solution is removed through the recovery well (Minerals

counsil of Australia, 2006: 195).

The biggest advantage of In-situ leaching is the reduction of capital and operating cost. The

reduction results in the economical exploit of low grade deposits which was not possible with

other leaching techniques. There is no need to remove, grind an mil the ore (Mudd, 2000:

528).

One important safety aspect required for In-situ leaching is the drilling of a monitor well. The

monitor well is used during and after mining to test for possible water contamination. The

confined deep aquifer must be allocated below the water table (Gupta,2003: 481).

In-situ leaching is an environmental friendly mining process which has minimal surface

impact (Minerals counsil of Australia, 2006: 195). The biggest environmental impact of In-

situ leaching is underground. The leaking of leaching agent into the underground water

table is frequent. The new research done by the IAEA on the environmental impact the acid

In-situ leaching done in the Soviet block is leading to a wide reviewing of In-situ uranium

mining (Mudd, 2000: 527).



Heap leaching

The next percolation leaching technique discussed is heap leaching. Heap leaching is

mainly done with waist rock dumps or low grade ore that contain valuable metals. Heap

leaching has been used since the eighteenth century (Gupta, 2003: 482). The life time of a

heap can be up to several years due to the slow dissolution kinetics. Figure 2.8 illustrates a

heap leaching process (Brugess, 2008):

Figure 2.8: Illustration of a heap leach process

The ore is crushed to a size smaller than 25 mm and pilled. The leaching agent is sprayed

at the top of the heap as show in Figure 2.8 and allowed to percolate through the heap. The

leaching agent dissolves the mineral and so it recovers the valuable metals (Gupta, 2003:

482). The pregnant solution is collected in a pond and then processed to retrieve the

valuable metals.

There is no need to for expensive milling and thus the main advantage of the heap leaching

is the reduction in capital and operating cost. There biggest disadvantage is the low

recoveries and the extended period for leaching (Gupta, 2003: 482).

The construction of a leaching pad is a major safety and environmental aspect. The

leaching pad should be impervious to the leaching liquid. The leaching pad acts as a barrier

between the environment and the heap. The leaching agent is a danger to the safety of the

workforce and environment.

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Atmospheric agitation leaching

The first agitation leaching technique discussed is atmospheric agitation leaching. The

leaching is done at atmospheric pressure and usually elevated temperature. The ore is in a

solution with the leaching liquid and is kept suspended, in contrast with the above

percolation techniques (Gupta, 2003: 483). Atmospheric agitation leaching is done in a

cylindrical steel tank, the steel is lined with rubber to protect the steel from corroding (Gupta,

2003: 483). The aggressive leaching conditions that is achieved in atmospheric agitation

leaching results in higher recoveries in shorter time compared to percolation leaching. The

slurry can be heated to the desired temperature using steam and the elevated temperature

will increase the reaction kinetics (Gupta, 2003: 483).

There are two methods to agitate the slurry namely pneumatic and mechanical. Mechanical

agitation is created using a propeller or a turbine mixer driven by a motor. Compressed air is

used for pneumatic agitation. The compressed air is introduced at the bottom of a cylindrical

vessel with a conical bottom.

There is an important safety and environmental hazard concerning the air agitation. The air

can transport poisons material into the environment from the leaching tanks. It is therefore

important to have an air monitoring system (Weber, 1996: 439). The environmental impact

is reduced because the leaching is done in a tank and the system is easily containable.

Pressure leaching

The next agitation leaching technique discussed is pressure leaching. Pressure leaching is

applied when the rate of dissolution is to slow for temperatures below 90C and results in the

recovery of the metal not to be economical. The technology is new but it has found

numerous applications around the world. The biggest drawback of pressure leaching is the

high capital, maintenance and operating cost (Minerals counsil of Australia, 2006: 185).

There exist a lot of advantages in the use of pressure leaching, the prevalent advantage is

the increase in dissolution kinetics. At the Aflease Uranium plant in Klerksdorp the

residence time is reduced from 18 hours to 2 hours (Bateman, 2004: 3).

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There are two types of pressure leaching. The first is pressure leaching without oxygen and

the second is with oxygen. Pressure leaching with oxygen is used for sulphide or uranium

ores. The oxygen is used as an oxidizing reagent in this system. The reaction is done in an

autoclave which is designed to handle the high temperature and pressure (Gupta, 2003:

484).

The slurry is pumped into the autoclave at high pressures. The slurry is then heated with

steam to achieve the desired temperature. Once the temperature is achieved the

exothermic reaction supplies the energy to maintain the temperature in the autoclave. In

some cases it is necessary to control the temperature with cooling water.

The autoclave is determined a pressure vessel and is therefore a great safety hazard if the

proper maintenance is not done on a routinely bases. The environmental impact is reduced

because the leaching is done in a vessel and the system is easily containable.

2.5.1.2. Leaching raw materials

The raw materials used for the leaching process is heavily dependent on the ore (Lurie et al.,

1962: 16). The one important ore characteristic is the numerous minerals present in the ore.

The amount of minerals that contains the valuable metal is important but the choice of

leaching agent is also highly dependent on the gangue minerals present in the ore (Merrit,

1971: 63). The leaching of ore can lead to the dissolution of other metals in the system

which causes impurities in the valuable metal product. It is important not to fully dissolve the

impurities in the ore, the lack of impurities in the leach solution will reduce the extraction cost

of the valuable metals (Lurie et al., 1962: 16). There are two main raw materials used in the

leaching process namely the leaching and oxidizing agent.

The two raw materials change the properties of the ore. The changes of properties are

necessary to ensure that the valuable minerals are dissolved. The properties that are

changed are the pH and the oxidation state of the metal in the mineral.

The pH is controlled using either an acid or an alkaline in which the metal dissolves. The

acid or alkaline is called the leaching agent. Sulphuric acid, nitric acid and hydrochloric acid

are the most widely used leaching agents in the industry. Sodium cyanide or sodium

hydroxide is used in alkaline leaching processes (Minerals counsil of Australia, 2006: 132).

The choice between acid and alkaline leaching is dictated by the ore used and the metal that

is extracted.



The nature of the host rock is very important. Acid leaching is used for silicone ores and

alkaline leaching is used in high lime content ore. The carbonate minerals will increase the

consumption of acid. The alkaline and acid react first due to the fast reaction kinetics. The

increase of acid consumption leads to greater raw material cost which can lead to the

uneconomical extraction of the metal (Lurie et al., 1962: 16).

The advantages of acid leaching are (Lurie et al., 1962: 16):

• Acid leaching is much more efficient than alkaline leaching

• The ore does not need to be ground to as fine a mess as with alkaline leaching and

this reduces the cost of milling.

The disadvantages of acid leaching are (Lurie et al., 1962: 16):

• The acids used for acid leaching is very corrosive and all the equipment that comes

in contact with the acid should be designed to withstand the corrosive thus increasing

capital cost.

• The acid should be neutralised before it can leave the system to reduce

environmental impact.

• Acid leaching is not very selective for Uranium and a lot of impurities is soluble in the

acid solution, this produces difficult separation of Uranium from the system.

The advantages of alkaline leaching are (Lurie et al., 1962: 17):

• Alkaline leaching is much more selective with regards to Uranium.

• The equipment does not need to be designed to resist acid corrosion and this

dramatically reduces the capital cost.

To explain the use of an oxidant in the leaching process, uranium leaching is used. Uranium

has two valence states in Uranium oxide, Uranium(IV) has a positive charge of 4 and form

the Uranium oxide complex namely UO2 and Uranium(VI) has a positive charge of 6 and

form the Uranium oxide complex namely UO3. U(VI) forms Uranium oxide in the tetravalent

state and has a molecular formula of UO2 and U(VI) forms Uranium oxide in the hexavalent

state that has a molecular formula of UO3 or UO2O. U(VI) is easily soluble in dilute non-

oxidizing acids such as HCl and H2SO4 where as U(IV) is not easily soluble in dilute non-

oxidizing acids. U(IV) is soluble in dilute oxidizing acids such as HNO3 (Venter &

Boylett,2009: 445).



For this reason it is very important to oxidize the U(IV) to U(VI) to reduce the cost of the

reactant used in the leaching process. There are a few common used oxidizing agents that

are easily produced, attained and economically viable for use, just to name a few; MnO2,

MnO2, NaCl, Fe

2(SO

4)3

and NaNO3. There is a down side to using the oxidation reactants,

side reactions will occur with other in-solute metallic complexes and reduce them to become

solute. This reduce the purity of Uranium in the solution and complicates the separation of

the Uranium from the system. The acid leaching efficiency is also dependent on the Fe2+

concentration. The Fe2+

is oxidized with the oxidation reagent to Fe3+

, the Fe3+

then acts as a

oxidation reagent and U(IV) is oxidized to U(VI) (Lurie et al., 1962: 18).

2.5.1.3. Leaching design parameters

Leaching is a chemical reaction and there is a lot of design parameters to keep in mind. The

effect that each design parameter has is different for each mining application. It is important

to try and design the leaching process keeping most of the parameters in mind. The

important design parameters for uranium extraction are as follows (Merrit, 1971: 63):

• The ore that is mined.

• The leaching agent concentration.

• The temperature of the reaction.

The leaching efficiency is dependent on the ore characteristics. The ore can contain several

different minerals that contain the metal that should be extracted. The dissolution of the

different minerals follows different mechanisms and thus the dissolution kinetics is different.

The different dissolution kinetics can result in the full dissolution of one mineral and only

partial dissolution of a different mineral (Lottering 2007, 1). One important example of this

phenomena is in the extraction of uranium. The maximum leaching efficiency of 90% is

achieved in the Vaal River, South Africa, and this is due to the different minerals present in

the ore. There are two predominant uranium minerals present in the Vaal River, uraninite

and brannerite. Uraninite is in abundance in the ore and is easily leached while brannerite is

the mineral that dissolves slowly. The kinetics of brannerite is so slow that economical

extraction of uranium from it is impossible.



Acid leaching is heavily dependent on the acid concentration, the requirement of the acid in

the leaching operation is to have enough free acid in the system to attack the Uranium

mineral inside the grounded to extract it, without excessive acid to dissolve the gangue

minerals. The reaction starts with the acid neutralizing the gangue lime minerals, in high lime

percentage ore the amount of acid consume to neutralize the lime minerals is as high as 200

kg per metric tone ore. After the Uranium is dissolved in the liquid, there should still be

enough acid present in the liquid to stop precipitation of the Uranium (Merrit, 1971: 63).

The time and temperature for the leaching process is interdependent of one another, when

the temperature is increased the reaction time is reduced dramatically. There exist an

optimum temperature and time for each type of ore; this optimization is compared to

operating cost and Uranium extraction. The optimum is usually designed for longer reaction

time and lower temperature (Lurie et al., 1962: 71). There are disadvantages with the

regards of increasing the temperature are as follows (Merrit, 1971: 71):

• The dissolution of gangue minerals will be much more, this will increase the

complexity of extracting Uranium,

• Increased corrosion on the equipment, even on certain stainless steel materials.

2.5.1.4. Leaching kinetics

A chemical reaction has occurred when a number of molecules from one or more species

are converted to form new species (Fogler 2006, 5). Reaction kinetics is the study on how

fast the reaction occurs (Fogler 2006, 4). The most useful scheme to classify chemical

reaction for an engineer is to classify the number of phases. The two classification schemes

are heterogeneous and homogeneous. One phase is present in homogeneous systems and

more than one phase is present for heterogeneous systems (Levenspiel 1999, 2). The

leaching process is a heterogeneous system because solid particles, the ore, is dissolved

with a leaching agent, a liquid.

There are numerous variables that affect the rate of a chemical reaction. The variables

range from easily manipulated variables, such as temperature and pressure, to complex

variables, such as mass and heat transfer. The rate of reaction is a single calculated

variable that takes into account all the variables that affect the rate of chemical reaction

(Levenspiel 1999, 3). The rate of reaction is discussed for the extraction of uranium.



The fact that more than one phase is present in the system, it is very important to

incorporate the mass transfer term into the usual chemical kinetics term. To easily calculate

the overall rate of reaction, the individual rate steps should be written on the same base

(Levenspiel 1999, 371). The base that is used with uranium leaching is interfacial surface of

the uranium particles. The design equation for the extraction of uranium is displayed in

Equation 2-1 (Ikeda et al., 1995: 267):

= −dC rSNdt V (2-1)

When the variables, S, N and V, stay constant for a system and only the leaching agent is

changed, the only change in the design equation is k, the dissolution rate of the uranium.

The variable k is the rate of reaction which is dependent on the leaching agent and the

process conditions. The equation of k for nitric acid as the leaching agent is as given in

Equation 2-2 (Ikeda, 1995: 267)

[ ]( ) − + 2.3

a b 3 3r= k k HNO NO (2-2)

The reaction rate is dependent on the concentration of the nitrate ion to the power of 2.3.

The value is reported to be 1 if the concentration of nitric acid is larger than 10 M (kinetic of

leaching 1994, 6). HNO2 is a very important reactant in the leaching of uranium using nitric

acid. HNO2 acts as an auto-catalyst (kinetic leaching, 4). It is not necessary to add HNO2

to the system because it is formed as a by-product in the leaching of uranium with nitric acid.

The variables ka and kb is temperature dependant and can be rewritten as Equation 2-3

(Ikeda,1995: 271):

(2-3)

The equation of r for sulphuric acid as the leaching agent is given as Equation 2-4 (Fleming,

1980):

(2-4)

4 795002.2 10 exp

368000.46exp

a

b

kRT

kRT

= × −

= −

[ ]1 1

8 3 22 32

613006 10 expr UO Fe FeRT

−− + + = × −



The important factors in the dissolution of uranium is the ferric and ferrous ions in the

system. The importance of the oxidation of iron is described in the above section. The

dissolution kinetics using nitric acid leaching is faster than the observed dissolution kinetics

of sulphuric acid leaching (Avvaru et al., 2007: 2109).

2.5.2. Ion exchange

The ion exchange process follows the leaching and clarification processes on a uranium

extraction plant. The leaching process is not very selective and produces leach liquor

containing many ion complexes and a low uranium concentration. The typical composition

of Rand leach liquor is reported by Pinkey et al. (1962:30) and shown in Table 2.4.

Table 2.4: Typical composition of Rand leach liquor

Ion complex Concentration Units

UO22+ 0.2 to 1.0 g/L

Free aid as H2SO4 3 to 7 g/L

Fe2+ 1 to 4 g/L

Fe3+ 1 to 4 g/L

Mn2+ 5 to 10 g/L

Silica soluble 1 to 2 g/L

SO42- , HS O4

- 20 to 40 g/L

NO3- 0 to 1.5 g/L

Cl- 0 to 0.6 g/L

SxO62- 0 to 0.06 g/L

Co(CN)62- 0 to 4 ppm

Further processes will purify and increase the concentration of the uranium complexes to

eventually effectively produce the uranium product. Ion exchange is very efficient in

selectively extracting uranium complexes from the leach liquor. In an ion exchange process,

anions or cations are adsorbed onto a solid particle, the resin bead, by replacing another ion

of a similar charge. Ion exchange is used in many processes including water softening

where calcium cations in the water replace sodium cations from a resin, thus reducing the

calcium in the water.



The application of ion exchange processes in the extraction of uranium was one of the most

significant developments in this process technology. Anionic ion exchangers are used to

selectively adsorb uranium complex anions. The uranium complex ions mainly depend on

the type of acid used in the leaching process. In the case where sulphuric acid is used as

leaching agent, uranyl sulphate complexes, such as UO2(SO4)22-, will form. The uranium in

the uranium complexes is in the U(VI) state.

Resins The uranyl complex anions are adsorbed onto a strong base anionic resin which has a high

affinity for uranium complexes. These resins are synthetic, organic-polymers based on

styrene with a cross-linking agent such as divinylbenzene. The degree of cross-linking

depends on the ratio between the styrene and the divinylbenzene in the resin. Ionic

functional groups are chemically added onto the resin to give it ion-exchange properties. In

strong base anionic resins these functional groups are usually quaternary ammonium, or

pyridinium groups for uranium recovery (Merritt, 1971:139). These functional groups have

mobile anions which can be replace by other anions. These ion exchange adsorbtion

reactions for uranyl sulphate complexes are generalized to Equation 2-5:

nRX + [UO2(NO3)n+2]n- → RnUO2(NO3)n+2 + nX- (2-5)

Here R represents the cationic functional group on the resin and X the mobile anion that is

replaced. These resins have different affinities to different ions in the system and bind more

tightly to some. This difference in affinities is quantified in the molar selectivity coefficient,

which describe the equilibrium state of the resin-solution system. It is reported by Merritt

(1971:140) that the affinities towards different monovalent anions for a typical strong base

anionic resin are decreasing in the order of Equation 2-6:

NO3

- > CN- > HSO4- > Cl- > HCO3

- > OH- > F- (2-6)

Resins have a higher selectivity coefficient towards ion with a higher valence. When the

valence of the mobile anion that is being replaced is lower than the adsorbed anion’s then a

lower concentration in the solution of the adsorbed anion will cause the equilibrium favours

the adsorption. Resins also have a higher affinity towards ions with the smallest solvated

volume, which in general decreases with the increase in atomic number. Uranium has a

very high atomic number, 92, which causes resins to have a very high affinity towards ion

complexes that contain uranium. There are however other ion complexes, such as

polythionates, cobalticyanides and molybdates, that will displace the uranium ion complexes

from the resin.



Depending on the degree of cross-linkage, the resin particle will swell to a certain size in

water. For resins operating in columns, more than 95% of the wet resin has a particle

diameter between 3.6 and 5.7 mm (20 – 50 mesh, U.S. standard screens). For resin in pulp

operation the resin beads are usually larger. Smaller resin beads has many structural and

kinetic advantages, but also causes a higher pressure drop over the column and is easily

lost through sieve plates. Merritt (1971:141) reports that in South Africa the pressure drop

over resin was recorded as approximately 22.6 kPa/m resin at a flow rate of 0.08 to 0.16 m/s

through the resin. Structural strength is also an important property for consideration with

resins, since resin attrition causes resin loss and new resin must be loaded from time to time

(Mirritt, 1971:138-142).

The maximum capacity of a resin is measured in equivalents per unit volume (wet) or mass

(dry). These equivalents are equal to the amount of moles of monovalent ions that is mobile

or can be replaced if only monovalent ions are loaded on the resin. Therefore if an anionic

resin has a capacity of one equivalent per gram, then one mole of hydroxyl ions can be

adsorbed per gram of resin. It should also be noted whether the capacity refers to wet or dry

resin. When the capacity is given per mass then it usually refers to dry resin mass and when

given per volume it usually refers to wet resin volume (Seader & Henley, 2006:557).

Adsorption

Since the kinetics of the adsorption is fast, equilibrium is reached very rapidly. When the

resin is loaded into a column with the pregnant leach liquor flowing slow enough through it

from above, the top fraction of the resin reaches equilibrium or saturation before any uranyl

complexes even start to leave the column. This results in the formation of an adsorption

front or section in which the adsorption of the uranyl complexes takes place. Above this

front the resin is saturated with uranyl complexes and below this front the resin has no uranyl

adsorbed because the solution passing over it is stripped from uranyl complexes. This

adsorption front is shown in Figure 2.9.



RE

SIN

Pregnant leach

liquor (Conc. = C)

Saturation

Resin

concentration R

esin bed height

Fresh resin

Saturated resin

Figure 2.9: Typical adsorption

As indicated in Figure 2.9, the adsorption front moves downwards through the column and

the thickness of the front depend on the kinetics of adsorption and the flow rate of the leach

liquor. When the lower end of the adsorption front reaches the bottom of the column it is

called “break through” and the first uranyl complexes starts to pass through. When the top

end of the adsorption front reaches the bottom of the column the column is saturated and the

exit and feed solution will have the same uranyl complex concentrations (Minerals Counsil of

Australia, 2006).

The pregnant leach liquor contains several ions as mentioned before. Anions such as

sulphate and hydrogen sulphate compete with the uranyl complex anions for adsorption onto

the resin. In leach liquor where nitric acid is during leaching, nitrate ions will be present and

competing for adsorption like the sulphate complexes. However these anions are necessary

in the pregnant leach liquor to ensure the equilibrium presence of uranyl anion complexes

such as UO2(SO4)22-. Therefore the optimal ratio of these competing anions and uranyl

complex anions should be determined to optimise the adsorption of uranium onto the resin.

It is reported by Merritt (1971, 147) that the uranium adsorption increases because of this

effect with an increase in pH, however at a pH value as low as 2.0 precipitation of uranium

will occur because of the presence of ions such as phosphate and arsenate. For this reason

the adsorption cycle is operated in a pH range of 1.5 to 2 (Merritt, 1971:147-148).



Although the affinity of the resin towards uranyl complexes are relatively high, there are

other ions that may form anion complexes that compete with uranyl complexes, such as

ferric and vanadic ions. Especially ferric anion complexes are present in greater

concentrations than the uranyl complexes. However, given sufficient time, complete

equilibrium is reached where a very small fraction of the resin is occupied by iron. Since the

kinetics of the adsorption is relatively fast, equilibrium is reached quickly leaving little uranyl

complexes to pass through the rest of the resin. This causes the iron to be adsorbed and

then later displaced by uranyl complexes. As a result of this effect an iron break through is

observed before an uranyl break through is observed (Gupta, 2003:548).

Elution

The adsorbed uranyl complexes are removed from the resin in the elution cycle. Here the

thermodynamic equilibrium is manipulated to favour the release of the uranyl complexes

from the resin. This can be done by introducing a solution or eluant with a high

concentration of a certain ion to replace the uranyl complexes, however this method is slow

and requires large volumes of eluant. A more efficient technique is to use ions in the eluant

that alter the uranyl anionic complex as well as displacing it. Dilute nitrate or chloride

solutions are used as eluant and produce non-adsorbable neutral or cationic uranyl

complexes. Sulphuric acid is also a possibility for use in an eluant. Table 2.5 shows these

eluant choices with suggested concentrations (Merritt, 1971:156-161).

Table 2.5: Suggested concentrations for eluant choices

Elution solution Concentration range (molar)

Chloride ion 0.5 – 1.5

Nitrate ion 0.8 – 1.2

Sulphuric acid 1.0

It is suggested by Merritt (1971:160) that the resin should be converted to the sulphate form

after elution with a nitrate or chloride ion solution is done. This will improve the adsorption of

new uranyl complexes.

Another elution process uses acidic ammonium nitrate (NH4NO3) as an eluant. This gives

the advantage that the adsorbed ferric complexes is eluted first and can be separated form

the rest of the elution with the higher uranyl complex concentration. The concentration of the

eluate, exiting the process, is shown for this system in Figure 2.10.



Figure 2.10: Eluate concentration profile

The x-section of the eluate seen in Figure 2.10 has a high ferric ion concentration and is sent

back to the pregnant leach liquor tank. The y-section of the eluate has a very high uranyl

complex concentration in comparison with the x-section and is sent to the next process. The

last z-section of the eluate is recycled to the eluant feed. This process is effective to remove

the ferric ions from the system and thus reducing the impurities in the system (Gupta,

2003:548-549).

Regeneration

Certain ions in the pregnant leach liquor have a very strong bond when adsorbed onto the

resin. These anion complexes may even replace uranyl complexes and if it is not removed

during elution, it is called poisoning of the resin. Examples are silica, molybdenum,

polythionates, cobalt cyanide complexes and thiocyanide. Basic fouling of the resin is also a

possibility.



Some of these poisons are temporary poisons which mean that that it is possible to remove

them from the resin through regeneration of the resin. The resins can be regenerated by

contacting it with a caustic soda solution. This removes the temporary poisons, but should

be done in stages to prevent resin damage through a pH shock. Other poisons are not

easily removed from the resins without damaging the resin itself and are called permanent

poisons. Therefore at some point the level of permanent poisoning will necessitate the

replacement of the resin (Merritt, 1971:163-167).

2.5.2.1. Ion exchange process alternatives

The previous sections described certain stages (adsorption, elution and regeneration) in the

ion exchange process. These fundamental stages along with other washing stages are all

incorporated in the processes used in practice. For this reason most of the existing ion

exchange processes are semi-continuous. In the section below the fixed bed column ion

exchange, moving bed column ion exchange and the resin in pulp processes is discussed

shortly.

Fixed bed column ion exchange

The operation of ion exchange through the use of fixed resin beds are commonly applied in

the uranium industry. Fixed resin beds are usually operated with three or four column

systems. Each of these columns goes though the discussed cycles of adsorption, elution

and back-washing. Regeneration of resin is not done as frequently and therefore the resin is

removed from the columns for this step.

These fixed bed columns are operated semi-continuously, which means that the pregnant

leach liquor is processed continuously, but the columns are not continuously in the

adsorption cycle. For the adsorption cycle, at least two fixed resin bed columns are

operated in series. The resin in the first column is saturated with uranyl complexes, which

occurs when the exit effluent has the same uranyl complex concentration as the pregnant

leach liquor fed to the column. This should take place before a breakthrough is observed in

the exit effluent of the second column, in which case uranium loss may occur. The exit

effluent from the adsorption cycle or the barren solution, which has a low pH value and

contains very little uranium, is sent back to the leaching section or an acid recovery plant

(Merritt, 1971:167,170).



As soon as the resin in the first column of the adsorption cycle is saturated with uranyl

complexes, the column is taken out of the adsorption cycle. Since there are some solid

particles from the pregnant leach liquor that accumulate on the resin, the column is back-

washed using slightly acidified water. If the solid particles are not removed, the pressure

drop over the column will increase and channelling may occur which will reduce the

efficiency of the resin. This back-wash step can also be done after the elution cycle, but

then the density of the resin is lower which limits the effectiveness of the back-washing.

During back-washing the resin bed may expand to twice its usual volume. The spent water

from the back-wash cycle is sent to the feeding tank of the ion exchange columns, since it

contains significant amounts of uranium (Weiss, 1985:24-27).

To remove the uranyl complexes, the saturated resin is next eluted in the elution cycle. The

use of elution medium is kept to a minimum by keeping the flow rate as low as possible

within the available time for elution. This also causes higher uranyl complex concentrations

in the eluate leaving the elution cycle. As mentioned before, split elution is used to increase

the uranyl complex concentration sent to the next process and may in some cases reduce

impurities such as ferric ions. It is suggested that the resin must be converted back to the

sulphate form to increase the recovery at the beginning of the adsorption cycle. This is done

by flushing the resin with a sulphate solution (Weiss, 1985:24-27).

The presence of resin poisons necessitates the regeneration of the resin. This regeneration

is usually done with caustic soda in another tank, therefore the resin is hydraulically pumped

out of the fixed bed column to a regeneration tank. It is important to prevent a physical

shock or “resin shock” when the resin is transferred from an acidic form to a hydroxyl from in

which it swells significantly. For this reason the resin is first contacted with 0.5% and then

5% caustic soda to remove the temporary poisons and regenerate the resin. Then the resin

is washed with neutral salt and water before converted back to the sulphate form (Weiss,

1985:24-28).

The efficiency of the resin decreases over time because of permanent poisons which is not

removed during regeneration. The capacity of the new resin should compensate for this

decrease and therefore columns with new resin have extra capacity. This extra capacity is

dependant on several factors such as rate of resin poisoning, regeneration frequency,

optimum resin life, variation in ore feed grade and plant amortization cost (Merritt, 1971:170).



The process equipment for fixed bed column ion exchangers usually consists of cylindrical

pressure vessels with slightly domed ends. A general ion exchange column has the

following dimensions, a diameter of 2m and a height of 3.7 m. It is usually constructed of

steel and lined with hard rubber on the inside to be acid proof and a variety of acid proof

piping is used. At the bottom of the vessel a diffuser piping system serves as an outlet for

the barren liquor and eluate and an inlet for the back-wash water. The resin fills about half

the vessel and lies on graded gravel or sand which supports it. During the backwash cycle

the resin volume expands approximately 100% to take up the whole column.

Figure 2.11 shows the layout for columns in an ion exchange system from Weiss (1985:24-

27). The columns have two piping inlet systems, one situated at the top of the column and

one in the middle. The inlet at the top of the column serves as an inlet for the pregnant

leach liquor and as an outlet for the back-wash water. The eluant is introduced into the

column right above the settled resin, in the middle of the column. During elution the space

above the resin in the column may be filled with either water or air, where the latter causes

least uranium loss due to diffusion of eluate into the water.

Figure 2.11 Fixed-bed ion exchange column layout



In Figure 2.11 shows the different valve to control the important flow cycles of the column

system. These valves are usually automatic controlled spring-closing, air-opening rubber-

lined diaphragm valves. Figure 2.11 also shows a manhole and windows in the columns,

which is necessary to check on the resin and for maintenance. The valve that allows the

hydraulic removal of the resin is seen near the bottom of the middle column in Figure 2.11.

The control of the fixed bed column ion exchange system is an important aspect for the

efficient performance of the process. The adsorption cycle determines the time available for

the other cycles. Since the loss of uranium must be prevented there must be an eluted resin

column available when a column on the adsorption cycle is saturated. When nitric acid is

used as eluant, the elution cycle is faster and only a total of three ion exchange columns are

necessary. For slower elution cycles, four ion exchange columns are needed to ensure the

continuous adsorption. These cycles are controlled automatically based on the time

required or volume flow required for each cycle. It is important that frequent checks are

done on effluent analysis to ensure that the process is optimally controlled and that no

uranium is lost (Merritt, 1971:172).

Moving bed column ion exchange

In the moving bed column ion exchange process the resin is moved between columns which

each perform a certain cycle similar to the cycles in the fixed bed system. A typical layout of

this system is two set of three columns for adsorption, one set of three columns for elution

and one column for back-washing. The resin is transferred hydraulically from one column to

the next.

There are three adsorption columns in series in the adsorption cycle, and the resin is

removed from the leading column when it is saturated. The adsorption cycle is then

operated with only two columns in series for as long as possible to obtain maximum loading

of the resin. When breakthrough is reached in the leading column, the third column is added

to the series, containing fresh eluted resin.



Saturated resin from the adsorption cycle is moved to a back-wash column where the resin

is back-washed before it is sent to the last column on the elution cycle. The elution is

different from the fixed bed system since here three columns in series are used of elution

instead of one. Since the fresh eluant is contacted with the first elution column containing

the resin whish has been partially eluted already, this system is a counter current system.

This gives the advantage of a high uranyl complex concentration in the eluate sent to the

next process (Merritt, 1971:174).

Each of the columns used in this process has an outlet for the transferring of the resin which

is usually located about 10 feet from the base. The typical adsorption and elution columns

have diameters of 2.5 m and a height of 4.5 m. The resin beds are nearly as deep as the

column heights since back-washing is not done in them. For this reason the back-wash

column is larger with a typical height of 5 m to accommodate the expanding resin bed during

back-washing.

Less piping is necessary since each column as a specific purpose and therefore the

connections is simpler. This also prevents the mixing of pregnant leach liquor and eluant

due to improper valve operation. The deeper rein beds also give the advantage of better

plant space utilisation and therefore less capital cost. As mentioned before, this system

produces an exit stream with a high uranyl complex concentration to the next process

because of multiple elution columns.

A general disadvantage of this system is the loss of uranium when the resin is not

transferred completely. The remaining resin from the adsorption cycle is loaded with uranyl

complexes which is lost when the column is reintroduced as the last column of the

adsorption cycle. This also decreases the total capacity of the resin.

Resin in pulp

There are several resin in pulp ion exchange processes. In this processes the leach slurry is

directly contacted with resin for the adsorption of uranyl complexes from the leach slurry. If

the solid particles in the leach slurry are small enough, the resin is easily separated from the

leach liquor with sieves. Therefore intensive solid liquid separation is not necessary, and

capital costs are reduced. In the following section the basket and continuous resin in pulp

processes are discussed.



Basket resin in pulp

In the basket resin in pulp process the resin is situated in baskets constructed of stainless

steel. The stainless steel frame is covered with stainless steel or plastic screen cloth with 5

mm holes. A group of these baskets are called a bank of baskets. This banks moves up

and down in rectangular tanks though which the leach slurry of eluant flows. The leach

slurry or eluant flows from one tank to the next through the use of air lifts of gravimetric flow.

Pumps are used to return leach slurry or eluant to the first tank, with piping from every tank

to the first tank and the next process.

These tanks or banks of resin baskets are operated in circular cycles. The leading tank in

the adsorption section is rinsed once saturated and then becomes the last bank in the

elution cycle. Once the leading bank in the elution cycle is saturated, it is again rinsed and

becomes the last bank in the adsorption cycle. The resin particles used in this system are

larger for effective separation from the leach slurry, 90% of the resin have particle diameters

between 6 mm and 8 mm. The larger particles require more time to reach saturation and

therefore a system of fourteen banks is usually used (Merritt, 1971:176-178).

Continuous resin in pulp ion exchange

In this process the resin and leach slurry or eluant flows in counter currents through a series

of tanks. The resin and leach slurry or eluant mixture that exits each tank is pumped or air

lifted to the top of the next tank. Here the resin is separated from the slurry with a vibrating

screen and falls into the tank below. The leach slurry or eluant from here goes to the tank

preceding the one it came from. This way the counter current flow is achieved. Resin traps

are situated at all exit streams to reduce resin loss to a minimum.

It is reported that six to eight stages are used in the adsorption section and seven to fourteen

stages are used in the elution section. This process requires more elution stages than seen

in basket resin in pulp systems because less efficiency is achieved due to lower eluant flows.

Higher flow rates are not introduced since it will cause low concentrations of uranyl

complexes in the eluate.

Finer resin is used in this operation with particle diameters between 6 mm and 4mm. More

resin attrition is observed because of the air agitation, screening and transport between

stages. A 20 % to 30 % resin loss per year is caused by attrition by this process. However it

uses less resin than the basket resin in pulp system (Merritt, 1971:178-180).



2.5.2.2. Ion exchange raw materials

In an ion exchange system there are a few reagent and product streams entering and

leaving the system. For the following discussion a fixed bed column ion exchange system is

used as base. In Table 2.6 a list of feed and products is given for this system as well as the

origin and destination of each.

Table 2.6: Feed and product streams for ion exchange

Feed streams

Name Description Purpose Origin/destination

Resin Strong base anionic type Adsorb uranyl

compexes Loaded periodically

Leach liquor Contains low concentration

uranyl complexes Feeds end product Clarifiers

Back-wash water Water Washes solid particles

out of resin Recycled water

Eluant Diluted sulphuric acid Remove uranyl

complexes from resin

Recycled from

solvent extraction,

return eluant

Flush water Water Rinse eluate from resin Recycled water

Product streams

Barren liquor Uranium removed Acid recovery

Back-wash water Water/solids Clarifiers

Eluate Uranium loaded diluted nitric

acid Carry uranyl complexes

Next process,

return eluate tank

Flush water Water Rinsed resin To return eluate

tank

As seen in Table 2.6, there are certain secondary systems necessary for the operation of

this section. A leach liquor feed tank, water recovery plant, return eluate tank and acid

recovery are required.



2.5.2.3. Ion exchange design parameters

In the design of a fixed bed column ion exchange process, the configuration and reagent

should be established first. There are certain structural parameters of importance for the

design of a column such as corrosion resistance and physical strength. The main factors

when sizing the column is the flow rates and time. For a given flow rate the time required for

the resin to reach saturation is dependant on the dimensions of the resin bed, the transport

kinetics and the equilibrium of the system. It is then further important to consider the time

required for the elution cycle in establishing the time for saturation of the resin in adsorption

cycle.

2.5.2.4. Ion exchange kinetics and equilibrium

The ion exchange process is written as a reaction in order to understand the equilibrium

behaviour. The reaction where anion Ay- replaces anion Bz- from the resin is written as

Equation 2-7:

( ) ( )z y y zy B zA z A yB− − − −+ = + (2-7)

Here the ion in brackets is the adsorbed ion while the other one is in solution. In the reaction

equation, y is the electron charge of the A ion and z the electron charge of the B ion.

The equilibrium for this reaction system is estimated with Equation 2-8:

z y

A BA,B y Z

B A

q cKq c

= (2-8)

Here q is the molar concentrations of the species in the adsorbed state and c the molar

concentration in the solution state. KA,B is the molar selectivity coefficient and is a constant

for dilute solutions and a certain ion pair and particular resin with a level of cross-linkage.

The molar concentration in the resin phase is taken as the number of equivalents per unit

bed volume or unit mass. In the solution phase the molar concentration is taken as the

equivalents per unit volume of solution.



In a system where the electron charge of the competing ions differs, the concentration of the

ion originally in solution has a great influence on the equilibrium. Manipulation of the

equilibrium equations shows that if the ion introduced in the solution has a higher electron

charge and a lower concentration, the equilibrium will favour the adsorption reaction (Seader

& Henley, 2006:565-567).

In a system with more than two competing ions, it may be assumed that the rest of the ions

do not affect the equilibrium of a certain pair. Then the system can be solved by satisfying

the different equilibrium equations for each pair of ion, and the mass balance of the system.

The equilibrium is important to show the available capacity of the resin.

According to Seader & Henley (2006:568) there are four steps that can influence the rate of

adsorption when looking at ion exchange. These four steps are:

• Liquid film diffusion (external mass transfer)

• Pore diffusion (internal mass transfer)

• Surface diffusion

• Chemical adsorption reaction

The step that is the slowest in a series of steps is assumed to be the rate determining step,

which is plausible if the specific rate is much slower than the rest. For ion exchange

systems only the external and internal mass transfer steps is considered to be rate

determining since the chemical reaction rate is fast. Either or both the external and internal

mass transfer steps can be rate determining. In general the external mass transfer step is

rate determining for systems with a low exchange-ion concentration, below 0.01N, and the

internal mass transfer step is rate determining for higher concentrations, above 0.1 N. In a

system with a very high selectivity for the adsorbing ion relative to other ions in the system it

has been observed that the external mass transfer step is rate determining. Another

observation is that divalent ions diffuse appreciably slower than monovalent ions through the

pores of the resin.

The kinetic behaviour of the internal mass transfer step is given by Equation 2-8.

(2-8)

2

e 2c 2 c qD

r r r t ∂ ∂ ∂

+ = ∂ ∂ ∂



In Equation 2-8 De is the effective diffusivity, r the distance from the centre of the resin bead

and t the time.

The kinetic behaviour for the external mass transfer step is given by Equation 2-9.

(2-9)

In Equation 2-9 the first term is the rate of molecular transport through the liquid film

surrounding the resin beads, kc is the transport coefficient, A is the area of the outer surface

of the resin bead, cbi is the concentration of component i in the bulk solution and csi is the

concentration of component i at the surface of the resin bead. It is important to note that the

transport coefficient is dependant on certain properties of the solution such as its flow rate

and viscosity. This dependence is shown in Equation 2-10.

(2-10)

Equation 2-10 shows that the transport coefficient kci of component i is dependant on the

diffusivity of component i in the mixture (Di), the resin bead diameter(Dp), the fluid mass

velocity (G), the viscosity of the solution (μ) and the density of the solution (ρ). In the

application of Equation 2-9, the concentration of component i at the surface of the resin bead

will be taken as the equilibrium concentration under the specific conditions. Therefore the

resin bead is taken as a pseudo particle with an even distribution of ions which is in a state

of equilibrium (Seader & Henley, 2006:568-572).

There are some short-cut methods which gives the kinetic profile of the adsorption cycle of

ion exchange in a fixed bed of resin. Figure 2.9 shows that there are two equilibrium phases

before and after the adsorption front. If the thickness and the downward velocity of the

adsorption front are known, then most of the design parameters can be determined. As

mentioned before the thickness and downward velocity of the adsorption front is dependant

on the flow rate of the pregnant leach liquor and the kinetics.

The Minerals Counsel of Australia (2006) gives such a short cut method to estimate an

approximate adsorption front thickness and downward velocity. This is given in Equations 2-

11 and 2-12.

( )i ic b sdN k A c cdt

= −

i

1/30.6pi

cp i

D GDk 2 1.1D D

µ = + µ ρ



(2-11)

In Equation 2-11 H is the thickness of the adsorption front, V is the solution throughput, C is

the feed concentration, A is the cross sectional area of the resin bed, qe is the capacity of the

resin for the specific component at equilibrium and e is the fraction of the bed which is void.

The downward velocity of the adsorption front, Ue, is given by Equation 2-12.

(2-12)

Here Ul is the superficial velocity of the solution over the resin beads.

The kinetics and equilibrium parameters of the systems are important to determine in order

to produce a viable design for an ion exchange process. The equations given in the section

above is not totally accurate, but it is the most accurate conceptual approaches to the

problems.

2.5.3. Solvent extraction

Solvent extraction is the commonly used name for liquid-liquid extraction. Perry (1997: 15-4)

gives a definitive definition for liquid-liquid extraction:

“Liquid-liquid extraction is a process for separating components in solution by their

distribution between two immiscible liquid phases.”

The principle of solvent extraction is the same as that of ion exchange, but instead of a solid

resin, a non-aqueous solvent is used. This process is mostly used in hydrometallurgy,

where an organic phase is contacted with the aqueous solution containing the desired metal

that needs to be purified (Gupta, 2003: 510).

There are generally three stages in the solvent extraction process; extraction, scrubbing and

stripping. These three stages are represented in Figure 2.12 (Gupta, 2003: 511).

( )eVCH q eCA

= +

( )l

ee

UCUq eC

=+



Figure 2.12: Typical solvent extraction process

The first stage is the extraction stage; here the metal is transferred from the aqueous

solution to the organic phase by bringing the two phases in contact. The products of the

extraction stage are the organic solvent loaded with the desired metal and the aqueous

extraction raffinate (Gupta, 2003: 510).

The organic product from the extraction stage is sent to the scrubbing stage where it is

treated with a fresh aqueous solution to remove any impurities present in the solvent. From

the scrubbing stage the organic solvent is then stripped to remove the solute. After the

stripping stage the organic solvent is regenerated and recycled back into the extraction

stage. The purpose of these three stages are to produce two aqueous solutions, one

containing most of the impurities, and the other most of the desired metal ions (Gupta, 2003:

510).

2.5.3.1. Process alternatives

The process through which solvent extraction takes place cannot change radically, there is

however different solvents and contactors that can be used.

The different solvents that can be used depend on the chemicals used in the previous

sections of the plant and the degree of separation required from the solvent extraction unit.

There are different processes concerned with extracting uranium complexes from sulphate

solutions. The two processes most commonly used are the Amex and Dapex processes,

where different extractant chemicals are used.



The main difference in solvent extraction processes, besides the composition of the solvent,

is the different types of extractors used. The different types of extractors that can be used

are (Madhavan, 2008):

• Mixer-settlers.

• Centrifugal devices.

• Column contactors (static).

• Column contactors (agitated).

There are different factors when selecting an extractor type; these factors are stage

requirements, fluid properties and operational considerations. Table 2.7 gives a comparison

between the different extractor types (Madhavan, 2008).

Table 2.7: Comparison of solvent extractor types

Property Mixer-settler Centrifugal extractor

Static columns Agitated columns

Number of

stages

Low Low Moderate High

Flow rate High Low Moderate Moderate

Interfacial

tension

Moderate to

high

Low to

moderate

Low to

moderate

Moderate to

high

Viscosity Low to high Low to

moderate

Low to

moderate

Low to high

Density

difference

Low to high Low to

moderate

Low to

moderate

Low to high

Floor space High Moderate Low Low

When choosing an extractor type, these properties need to be taken into account, and then

an extractor type can be chosen for the desired process.

2.5.3.2. Raw materials added

The raw materials added to the solvent extraction unit are the organic compounds that make

up the organic solvent, the solvent components consist of the following (Gupta, 2003: 512):



• Extractant.

• Diluents.

• Modifier.

The extractant is the active component in the solvent and is responsible for extracting the

metal from the aqueous phase. The extractant needed for a sulphate system is an anion

exchanger. The most common anionic extractant is long-chain alkylaminies (alamines),

alamines can be classified as being of primary (RNH2), secondary (R2NH), and tertiary (R3N)

nature. Tertiary alamine is used for the sulphate system (Gupta, 2003: 513).

The diluent is the carrier component in the solvent, and has the purpose of lowering the

viscosity of the extractant to simplify transport. The widest used diluent is kerosene, due to

its low cost and high flash point. During the extracting phase, there may form a third phase

when the organic phase splits into two. To solve this problem, a modifier is added. For the

purpose of this project, isodecanol (C10H22O) is used as the modifier added to the solvent

(Gupta, 2003: 515).

A summary of the components in the organic solvent are:

• Tertiary alkyl amine (Alamine® 336).

• Kerosene.

• Isodecanol (C10H22O).

To design a solvent extraction section and to decide which solvents to use, there are certain

design parameters that needs to be considered, these parameters are discussed in the next

section.

2.5.3.3. Design parameters

There is one function of a solvent extraction unit, and that is to mix two liquid phases, form

and maintain droplets of dispersed phase and later separate the phases again. The two

design parameters that are the most important are mixing and settling.



Both mixing and settling is determined by the physical properties of the phases present in

the system. The amount of mixing is very important, if too little mixing is provided large

droplets will form and decreases the interfacial area which reduces mass transfer and

decrease stage efficiency. More mixing will minimize mass transfer resistance during

reactions and extraction (Madhavan, 2008).

Apart from the physical properties, settling is determined by the amount of mixing. The most

problems in settling occur when the phases are mixed too much, this forms an emulsion

which is difficult to separate. If an emulsion is formed, it needs to be settled over an

extensive period of time (Madhavan, 2008).

The other design parameters present in solvent extraction are the selection of the

components in the organic solvent used to extract the metal from the aqueous solution and

the pH of each stage in the solvent extraction process.

2.5.3.4. Kinetics

The two sections where kinetics plays an important role are the extraction and stripping

stages. The solvent extraction of uranyl sulphate, by using tertiary amines as extractant, is

characterized as having extremely rapid reaction kinetics for both the extraction and

stripping stages. Due to this rapid kinetics short residence times is expected in the mixers,

the residence times range between 45 seconds and 2 minute (Kordosky et al.: 10)

2.5.4. Precipitation

The precipitation of uranium is one of the final steps in producing the ADU product. This

process is necessary to obtain a solid product that meets certain grade and purity

specifications. These specifications are important as there are penalty schedules in place,

which means that the revenue will decrease if the specifications are not met.



Precipitation of metal ions from aqueous solution is widely used in hydrometallurgical

processes as a means of separation, purification and disposal. The precipitate is formed by

changing solution pH, solvent concentration, solution temperature, or by the addition of a

reagent to the solution, also known as reactive crystallization. By adding a reagent to a

metal ion-containing solution, a compound will form which has a very low solubility and so it

readily precipitates. If the reagent is something other than water, it is known as ionic

precipitation and the added reagent functions by contributing anionic species. The

interaction of these anions with the cationic metal in the solution forms the compound

(Gupta, 2003:535).

The precipitation of uranium diuranate (ADU) is largely depended on the purity and

composition of the uranyl solution entering the process. If the solution is a product of ion

exchange, the iron concentration might be too high and an additional precipitation step is

introduced to precipitate the iron and other impurities. It is also important to note that the

ammonium diuranate (ADU) that forms has a very high concentration of uranium. This

uranium is in a complex form which means that it is dissolvable and can therefore be

absorbed into the body. This will cause heavy metal poisoning. Special care must be taken

when working with this process to make sure that there is no leakage of this poisonous

precipitate to the environment through the air.

Manipulating the techniques and conditions of the precipitation process, the selectivity and

physical characteristics of the precipitate can be controlled. There are several alternatives

for precipitation, which all have their specific application (Merritt, 1971:240).

2.5.4.1. Precipitation process alternatives

The product of precipitation, ADU, is ultimately converted to UO2 powder for the use in

nuclear power plants. It is thus important to obtain a precipitation product with satisfactory

physical characteristics, such as settling and filtering characteristics, and which also meet all

required product specifications for uranium and impurities content (Merritt, 1971:240).

There are two common precipitation methods used, the first involving direct neutralization

with a base such as lime or ammonia, and the second, direct precipitation from acid solution

with hydrogen peroxide (Merritt, 1971: 240-247).



Neutralization

The neutralization process used for precipitation is involved with the lowing of the pH of the

homogeneous solution to the point at which the metal complexes begin to break down and

cations are released from the soluble complexes. This is done in solutions containing the

desired anions to precipitate the product. Different reagents can be used to decrease the pH

of the solution, depending on the desired precipitate (Cartwright et al, 1967:667).

The precipitation of uranium will take place at pH numbers of 6.5 to 8, but the uranium

solution entering the precipitation process usually has a pH level of approximately 1.4 to 4.

This means that the solution must be neutralized to this level and can be done using caustic

soda, magnesia or ammonia. The reagent most widely used is gaseous ammonia, which

can be introduced into the uranyl sulphate solution and will cause the precipitation of

uranium in the form of ammonium diuranate (ADU) according to the overall reaction:

2 4 4 4 2 2 7 4 2 4 22UO SO + 6NH OH = (NH ) U O + 4(NH ) SO + 3H O

The characteristics of the uranium precipitate are dependent on the conditions maintained

during the precipitation process, including temperature, pH, feed solutions and the rate of

precipitation.

Careful pH control is necessary to ensure that a dense and readily filterable precipitate is

produced. A general method for this type of precipitation is to increase the pH value of the

solution gradually in separate tanks as a continuous process, which means that the solution

is pump form tank to tank. This causes a crystalline precipitate to form, which is desired

because this precipitate is easy to handle and wash (Merritt, 1971: 240 – 246).

Precipitation with hydrogen peroxide

A second division of cation release precipitation method is the release of cations at a

constant pH. This method involves the adjustment of the pH of the complex, and then the

complex is slowly destroyed by boiling with hydrogen peroxide to produce a dense

precipitate (Cartwright et al, 1967:667).



This method can be applied to precipitate uranium as the complex from which it is

precipitated is stable, the uranium precipitate is insoluble and catalytic breakdown of

hydrogen peroxide is possible by the precipitate particles. The precipitation process to

precipitate uranium is carried out by adding the hydrogen peroxide as an aqueous solution

slowly to the uranyl sulphate solution. The precipitate that forms is uranium peroxide,

UO4●xH2O which is a crystalline product that is easily handled. The reaction is shown below

(Brown, 1982):

++ → +2+

2 2 2 4UO H O UO 2H

The temperature and pH level during this process must be controlled to achieve the desired

product characteristics. The temperature is kept between 30̊ C and 65˚C and the pH is kept

at 2.8 for the optimum formation of a crystalline product with minimum contaminant.

The cost for using this process is somewhat higher than the conventional method of

neutralization because the raw materials are more expensive. But a higher purity product is

obtained which might save costs if penalty charges are imposed (Merritt, 1971: 247 – 248).

2.5.4.2. Precipitation raw materials

A two stage precipitation process is usually used, the first stage to precipitate impurities such

as iron, aluminium, titanium and thorium and the second stage to adjust the pH to a value

that will precipitate the desired product.

For the first stage lime is used to precipitate the impurities, which is also an economical

advantage since the quantity of the other, more expensive, reagents needed for

neutralization is decreased.

The raw materials for the second stage are dependent on the method chosen for

precipitation. If the neutralization method is chosen the reagents necessary can be caustic

acid, magnesia or ammonia. The total consumption of the neutralizing reagents will range

from 0.09 to 0.18 kg per kg of U3O8 in the precipitate. The choice between these three

reagents depends on several factors shown in Table 2.8.



Table2.8: Comparison of neutralization reagents

Cost Product characteristics

Reaction speed

Ease of control

Product contamination

Caustic soda

R 4 200/ton Slimy product,

difficult to

handle

Satisfactory Suitable for

automatic

control

Exceed sodium

specifications

Magnesia R 5 000/ton More crystalline

product

Slow Not suitable for

automatic

control

Below

specifications

Ammonia R 15 000/ton Slimy product Satisfactory Suitable for

automatic

control

Below

specifications

If caustic soda is used, it is usually introduced into the process as a 10% NaOH solution. If

ammonia is used, it is vaporized to a gas by heating just prior to use and as it is important to

have good dispersion of the gas, the ammonia is mixed with 2 to 4 parts of air before

addition.

Considering the precipitation process using hydrogen peroxide, a suitable base such as

ammonium hydroxide is also added to the solution to maintain the desired pH by neutralizing

the acid formed in the precipitation reaction. The 30% hydrogen peroxide solution is added

in at least stoichiometric quantity, but if the peroxide is in excess the precipitation is more

likely to be complete. The neutralization of the acid formed also requires a minimum of the

stroichiometric quantity of the suitable base.

The precipitation process using urea (also known as carbamide, (NH2)2CO) only needs this

one reagent. About a kilogram of urea is added to a kilogram of uranium as a filtered

solution which generates enough ammonia to precipitate the desired product (Merritt, 1971:

240 – 246).



2.5.4.3. Precipitation design parameters

There are two major considerations that have to be kept in mind when evaluating the product

that is produced by precipitation. The first consideration is the grade and recovery of the

product and the second is the physical characteristics which make the product easy to

handle and wash. These factors can be controlled by certain design parameters which

include the pH of the solution, the addition of reagent and the temperature at which

precipitation takes place.

pH The chemical and physical properties of the precipitate product are affected by the changes

of pH of the precipitation solution. The size of ADU crystallites and agglomerates is

decreased with an increase of solution pH. As the agglomerate decrease the filterability of

the ADU slurries also decrease (Janov et al, 1971:1). It has been reported that fine

sediments can be produced at pH above 6, which yields a high-density pellet, favorable for

nuclear energy plants. But at pH values below 5 coarse sediments are produced that settle

rapidly and are easily filtered (Wilson, 1996:146). In practice, different procedures are used

to control the pH at which precipitation should take place. A general procedure is to

increase the pH value gradually in separate tanks as a continuous process (Merritt,

1971:242).

Addition of reagent

When precipitation of a product is done by adding a reagent to the solution, a reaction takes

place to produce another chemical that is almost insoluble in the resulting solution. This

reaction produces a large degree of supersaturation, which is a very important factor in the

precipitation process. The extent of supersaturation is dependent on the ionic

concentrations in the reagent and solution before mixing. This supersaturation causes

primary nucleation, which are very small particles that are formed with a crystalline structure.

The extent of supersaturation influence the time it takes for precipitation to begin, the

number of particles formed per unit volume solution and the particle growth rate (Seader &

Henley, 2006:671-672).



Temperature

The temperature at which precipitation takes places has an influence on the rate of

precipitation settling and the precipitation time. The precipitation time shows an exponential

dependence on the temperature of precipitation and thus the time can be optimized to

reduce the quantity of precipitant but still maintaining the product specifications. Low

precipitation rates produce a product of a certain required sinterability. At these low

precipitation rates low temperatures favours the dispersion of precipitate particles while high

temperatures favour the agglomeration of the precipitate, which is the desired effect (Murty

et al, 2001).

A study of the temperature effects on the precipitation process using hydrogen peroxide

showed that the grade and recovery of the product decreased with an increase in

temperature, because of the faster decomposition of hydrogen peroxide with increase in

temperature (Gupta et al, 2004).

2.5.4.4. Precipitation kinetics

The kinetics of the precipitation process is complex, involving nucleation and crystal growth,

with supersaturation as the driving force for these steps. Precipitation occurs when two

reacting solutions forms a solid product with low solubility. This solid product is formed by

fast crystallization resulting in large numbers of very small crystals.

Precipitation will only take place a certain time after the development of supersaturation,

called the induction period, because of the slow growth of small particles. When the

supersaturated concentration of the solute is high, spontaneous nucleation of very small

crystals occur, lowing the concentration of the solute. As the concentration of the solute

decreases a metastable region is reached where crystals can grow but cannot nucleate, thus

if no crystals are present, none can be formed. At a certain point, equilibrium is reached

between the saturated solution and crystals that is formed. If the concentration of the solute

further decreases, the solution will be unsaturated and crystals of all sizes will dissolve.



As mentioned, crystallization consists out of two steps, nucleation and crystal growth.

Nucleation is the formation (birth) of crystals and can be primary or secondary. Secondary

nucleation is mainly observed in industrial crystallizers where crystalline surfaces are

present and large crystals are desirable. The mechanism which is typically encountered in

industrial applications for secondary nucleation is referred to as contact nucleation and

occurs when crystal collide with each other and the crystals collide with metal surfaces such

as the vessel wall or agitator blades. An empirical power-law function is widely used to

describe the secondary nucleation and is given in Equation 2-13 (Seader & Henley, 2006:

659).

=0 b j rN TB k s M N (2-13)

In Equation 2-9, B0 is the rate of homogeneous primary nucleation (number of nuclei), s is

the relative supersaturation, MT is the mass of crystals per volume of magma, N is the

agitation rate and the constants kN,b,j, and r are determined from experimental data.

The crystal growth can be explained by a two-step theory, referred to as the diffusion-

reaction theory. The first step of which is the mass transfer of solute from the bulk solution

to the crystal-solution interface and the given by Equation 2-14 (Seader & Henley, 2006:

659).

( )= −c idm k A c cdt (2-14)

With: dm/dt = rate of mass deposited on the crystal surface

A = surface area of the crystal

kc = mass-transfer coefficient

c = mass solute concentration in the bulk supersaturated solution

ci = supersaturated concentration at the interface

The second step is assumed to occur at the crystal-solution interface, in which solute

molecules are integrated into the crystal-lattice structure and this step is given in Equation 2-

15 (Seader and Henley, 2006: 659).

( )= −i sdm k A c cdt (2-15)



With: cs = Mass solute concentration in the solution at saturation

ki = Kinetic coefficient

The combined kinetics of these two-step theory is given by Equation 2-16 (Seader and

Henley, 2006: 659).

( )−=

+s

c i

A c cdm1 1dt k k

(2-16)

The mass-transfer coefficient will be depended on the velocity of the solution, which means

at low velocities the growth rate will be controlled by the first step. The second step will be

important if the mass-transfer coefficient is large compared to the kinetic coefficient (Seader

& Henley, 2006: 658-660).

The growth rate of particles can also influence the number of particles formed. If rapid

growth is observed, co-precipitation may occur which make it difficult to obtain a pure

precipitate product. The growth rate can be controlled by the mass transfer of the ions to the

particle surface and/or integration of ions into the particle crystalline structure (Seader &

Henley, 2006: 671-672).

Using the kinetics supplied above, the volume and residence time of a precipitation vessel

can be calculated using the rate at which nucleation and crystal growth occurs.

2.6. Economic evaluation

The last few years showed a significant interest in the uranium market, encouraged by rapid

raising prices. The growing uranium market resulted in various uranium mining projects,

including new mining projects and expansion projects. A uranium project survey was done

in 2007 to investigate a few of these uranium projects. This survey provides a good

indication of capital and operating costs for these various plants and a summary of this

survey is given in Table 2.9.



Table 2.9: Summary of uranium project survey (Damarupurshad, 2007: 6-7)

Project Company Capital costs (R million)

Operating costs

Dominion uranium SXR Uranium One 1 036.64 R 98.89/lb of U3O8

Ezulwini uranium &

gold

First Uranium

Corporation

3 219 R 419.58/ton ore

milled

Buffelsfontein

uranium & gold

First Uranium

Corporation

2 486.4 R 18.28/ ton ore

milled

The Ezulwini uranium and gold project investment and cost are explored in more detail. The

initial capital investment consists out of the mine life capital, including contingency, which is

R 2 072 million and the pre production capital which is R 1 147 million and is expended over

three and a half years. The average operating cost over 19 years is R 419.58 per ton of ore

milled. The net positive value for this project was calculated to be R 1 909.2 million using a

gold price of $ 500/oz and a uranium price of $ 40/lb with a rand to dollar exchange rate of R

7.40. The underground mining capacity of the Ezulwini project is up to 200 000 tons per

month and it was expected that at the end of 2008 the uranium production will reach an

output of 130 000 tons/month (Damarupurshad, 2007: 7).

Considering the economic evaluation of the Ezulwini project, it can be seen that a uranium

extraction project is feasible, yield a high net positive value (NPV). This high NPV is also an

indication that the capital investment can be sold at the end of the project lifetime to earn

back the money invested.


57 Chapter 3: Process development

Chapter 3: Process development

The proposed project requires the upgrade of a uranium extraction plant. In order to meet

the requirements given for this project it is necessary to understand the processes that are

used to extract the uranium. These processes are best understood if a certain design

procedure is followed to develop the process units. This section provides a detailed design

procedure that is given by Douglas (1988) which follows a systematic approach using certain

hierarchy of decisions listed below.

• Level 0: Input information

• Level 1: Batch versus continuous

• Level 2: Input-output structure of the flowsheet

• Level 3: Recycle structure of the flowsheet

• Level 4: Separation system

o 4a: Vapor recovery

o 4b: Liquid recovery

o 4c: Solid recovery

• Level 5: Heat integration

Each of these decision levels will now be discussed in order to understand the different

processes necessary to extract uranium. These levels will also allow continuous economic

evaluations during the design procedure to test economic feasibility. After these levels of

design are discussed, a detailed process description is given.

3.1. Level 0: Input information

The basic information required for the development of a process is given in the Level 0 of the

Douglas design procedure. Without the information gathered for Level 0, it is impossible to

start a conceptual design.



3.1.1. Reactions and reaction conditions

The extraction of uranium usually requires four sections with each section consisting of their

own reactions and operating conditions. The most important operating conditions are the

temperature and pH. The extraction of uranium consists of leaching, counter-current

decantation, ion exchange, solvent extraction and precipitation sections. Leaching is the first

section in the extraction of uranium and is considered the most important. The operating

conditions for leaching are a pH of 2 and a temperature of 30 °C throughout the leaching

section.

Table 3.1: Leaching reactions

Reaction Stoigiometry

Sulphuric acid

dissociation H2SO4 + H2O → HSO4

- + H3O+

Sulphuric acid

dissociation HSO4

- + H2O → SO42- + H3O+

Nitric acid

dissociation HNO3 + H2O → NO3

- + H3O+

Muscovite leaching KAl3Si3O10(OH)2 + 10H3O+ → 16H2O + K2+ + 3AL3+ + 3SiO2

Chlorite leaching Mg2Al4Fe2Si2O10(OH)8 + 20H3O+ → 2Mg2++ 2Fe2+ + 4Al3+ + 2SiO2 +

34H2O

Pyrophylite

leaching Al2Si4O10(OH)2 + 6H3O+ → 2Al3+ + 4SiO2 + 10H2O

Pyrite leaching 6FeS2 + 30HNO3 + 3SO42- → 6Fe(SO4)2

- + 3H2SO4 + 30NO + 12H2O

Albite leaching NaAlSi3O8 + 4H3O+ → Na+ + 2Al3+ + 3SiO2 + 6H2O

Uraninite leaching UO2 + 2NO3- + 4H3O+ → UO2

2+ + 2NO2 + 6H2O

Brannerite

leaching UTi2O6 + 2NO3

- + 4H3O+ → UO22+ + 2NO2 + 6H2O + 2TiO2

U-phosphate

leaching UO2ClPO4 + 2NO3

- + 4H3O+ → UO22+ + 2NO2 + 6H2O + Cl- + PO4

2-

Coffinite leaching UO2·2H2O + 2NO3- + 4H3O+ → UO2

2+ + 2NO2 + 8H2O

Coffinite leaching UO2SiO2 + 2NO3- + 4 H3O+ → UO2

2+ + 2NO2 + 6H2O + SIO2

Uraninite leaching UO2 + 4NO2 + 2H3O+ → 2HNO3 + UO22+ + 2NO + 2H2O



Uranyl sulphate

complex formation UO2

2+ + 3SO42- → UO2(SO4)3

4-

Silicon dioxide

dissolve SiO2 + 2H2O → SiO2·2H2O

The second section in the extraction of uranium is the ion exchange process. In the ion

exchange section uranyl sulphate complexes are selectively adsorbed onto a resin and

desorbed again; this allows separation and upgrade of concentration. Table 3.2 contains the

reactions, and reaction conditions for the ion exchange section.

Table 3.2: Reaction information for ion exchange

Reaction Stoigiometry Temperature

(°C) pH

Uranyl sulphate


2+ + 3SO42- → UO2(SO4)2

4- 25 2

Iron sulphate

complex formation Fe3+ + 2SO4

2- → Fe(SO4)2- 25 2

Uranyl sulphate

adsorption 2R2-SO4 + UO2(SO4)3

4- → R4-UO2(SO4)3 + 2SO42- 25 2

Iron sulphate

adsorption R2-SO4 + 2Fe(SO4)2

- → 2R-Fe(SO4)2 + SO42- 25 2

Nitric acid

dissociation HNO3 + H2O → NO3

- + H3O+ 25 2

Uranyl sulphate

elution R4-UO2(SO4)3 + 2SO4

2- → 2R2-SO4 + UO2(SO4)34- 25 1.5

Iron sulphate

elution 2R-Fe(SO4)2 + SO4

2- → R2-SO4 + 2Fe(SO4)2- 25 1.5

Silicon dioxide

precipitation SiO2·2H2O → SiO2 + 2H2O 25 1.5

Silicon dioxide

dissolve SiO2 + 2H2O → SiO2·2H2O 25 1.5

The next section is the solvent extraction section; here the uranyl sulphate complexes are

removed from the aqueous phase into an organic phase. The up-concentration of uranyl

sulphate complexes is then done by stripping the complexes from the organic phase into a



clean aqueous phase. Table 3.3 contains the reactions, and reaction conditions for the

solvent extraction section.

Table 3.3: Reaction information for solvent extraction

Reaction Stoigiometry Temperature

(°C) pH

Sulphuric acid

dissosiation H2SO4 + H2O → HSO4

- + H3O+ 25 1.5

Sulfuric acid

dissosiation HSO4- + H2O → SO4

2- + H3O+ 25 1.5

Uranyl sulphate


2+ + 3SO42- → UO2(SO4)3

4- 25 1.5

Solvent

preparation 2R3-N + H2SO4 → (R3NH)2SO4 25

Solvent extraction UO2(SO4)34- + 2(R3NH)2SO4 → (R3NH)4UO2(SO4)3 + 2SO4

2- 25

Solvent stripping (R3NH)4UO2(SO4)3 + (NH4)2SO4 → (R3NH)2SO4 +

(NH4)2UO2(SO4)2 + (NH4)2SO4 25 2.5 to7.5

The last section in the extraction process is precipitation; here the pH is manipulated to

precipitate a solid ADU product. The solid precipitate allows for more economical separation

of the product which is sent to NUFCOR which should comply to external specifications.

The reaction conditions for the precipitation are a pH of 7.5 and a temperature of 30 °C. The

reaction for the precipitation is given in Equation 3-1:

2UO2SO4 + 6NH4OH → (NH4)2U2O7 + 2(NH4)2SO4 + 3H2O (3-1)

It is important to know the kinetics and conversion of each reaction which is needed to size

the chosen processing equipment. The lack of information for the kinetics of mineral

processes limits the design process, and therefore assumptions are made. Most of these

assumptions are made from literature and is based on the conversion of reaction. The

available reaction kinetics and conversion assumptions are given in Table 3.4.



Table 3.4: Reactions kinetics and conversion

Reaction Reaction rate Conversion Source

Sulphuric acid

dissociation Equilibrium 100% of H2SO4 -

Nitric acid dissociation Equilibrium 100% of HNO3 -

Muscovite leaching - 6.2% of Muscovite Snäll & Liljefors

Chlorite leaching - 60% of Chlorite Snäll & Liljefors

Pyrophylite leaching - 5% of Pyrophyllite Lottering et al.

Pyrite leaching - 3% of Pyrite Karaca et al.

Albite leaching - 1% of Albite Snäll & Liljefors

Uraninite leaching [ ] − − − × +

2.342 3

79500 368002.2 10 exp 0.46exp HNO NORT RT

99.99% of Uraninite Zhao & Chen

Brannerite leaching - 10% of Brannerite -

U-phosphate leaching - 1% of U-phosphate -

Coffinite leaching - 90% of Coffinite Lottering et al.

Uranyl sulphate

complex formation - 100% of UO2

2+ -

Silicon dioxide dissolve - 0.001% of SiO2 -

Iron sulphate complex

formation - 100% of Fe3+ -

Uranyl sulphate

adsorption -

40% of Resin

capacity Merrit et al.

Iron sulphate adsorption - 1% of Resin


Sulphate adsorption - 50% of Resin


Uranyl sulphate elution - 100% of Uranyl

sulphate adsorbed -

Iron sulphate elution - 100% of Iron

sulphate adsorbed -

Sulphate elution - 100% of Sulfate

adsorbed -

Silicon dioxide

precipitation -

100% of dissolved

Silicon dioxide -

Solvent preparation - 100% of R3-N Rydberg et al.



Solvent extraction Equilibrium 99.99% of Uranyl

sulphate complex Rydberg et al.

Solvent stripping Equilibrium 99.99% of loaded

solvent Rydberg et al.

Precipitation - 98.49% of Uranyl

sulphate complex Mellah et al.

The values from Table 3.4 are used for the mass and energy balance, process design and

detail design.

3.1.2. Desired production rate and purity

The plant is designed to process 360 000 ton ore per month, at a uranium recovery of 78%.

The desired product is ADU in a slurry phase which consist of 35 mass% U3O8. Penalties

are paid for product which does not meet product specifications. Using the above

information, the amount of ADU produced per month is approximately 89 ton.

3.1.3. Raw materials

As stated in Douglas (1988: 104), laboratory studies are usually carried out with pure

chemical reagents. However, this is not the case for industrial applications, thus it is

important to know which impurities are present. The impurities can cause unwanted side

reactions that should be monitored.

The raw materials needed for leaching includes the ore feed, nitric acid, sulphuric acid,

Magnafloc 90L and potable water. Concentrated nitric acid with a 68 mass% is used which

is a standard product purity from a distillation column, due to an azeotrope formed with

water. A 98 mass% sulphuric acid is used which is already available on the existing plant.

Magnafloc 90L is used as a clarifier in the counter-current decantation section. Potable

water is used to dilute the acids and to achieve a slurry with a SG of 1.6 for the ore feed.

The ore composition is given in Table 3.5 (Lottering et al., 2007:18).



Table 3.5: Ore composition

Mineral Mass % Quartz (SiO2) 70.200

Muscovite (KAl3Si3O10(OH)2) 10.100

Chlorite (Mg2Al4Fe2Si2O10(OH)8) 2.000

Pyrophylite (Al2Si4O10(OH)2) 9.700

Pyrite (FeS2) 1.300

Albite (NaAlSi3O8) 4.800

Uraninite (UO2) 0.014

Brannerite (UTi2O6) 0.013

U-phosphate (UO2ClPO4) 0.001

Coffinite (U(SiO)41-x(OH)4x) 0.002

Resin, dilute sulphuric acid, caustic soda and potable water are the raw materials used for

the ion exchange section. The resin used is a strong base anion exchange resin, Ambersep

TM 400 SO4 which has quaternary ammonium as functional group. Due to resin poisoning,

caustic soda is used to regenerate the resin. The caustic soda is received in the aqueous

phase and potable water is used to wash the solid impurities from the resin and to dilute

concentrated sulphuric acid.

The raw materials needed for the solvent extraction section are kerosene, isodecanol, alkyl

amine and demineralised water. Kerosene is a collection of organic substances which are

inert to the system, and only acts as a carrier fluid. Isodecanol and alkyl amine is pure

organic chemical substances acting as the third phase modifier and extractant respectively.

Precipitation raw materials include ammonia gas and demineralised water. A summary of all

the raw materials used is given in Table 3.6.



Table 3.6: Summary of raw materials

Raw material Purity Description Price

Alamine (Alkyl amine) Pure Solvent extractant R 57 260 per ton

Ammonia Pure gas Precipitation reactant R 15 000 per ton

Calcium oxide Pure Neutralization R 139 per ton

Isodecanol Pure Solvent modifier R 22 410 per ton

Kerosene Pure Solvent diluent R 5 000 per m3

Magnafloc 90L Pure Clarifying agent R 22 500 per ton

Nitric acid 68 mass% HNO3 Oxidation agent R 1 500 per ton

Potable water Municipal Multiple uses R 7 per m3

Resin (Ambersep TM 400) Pure Ion exchange resin R 57 000 per ton

Sodium carbonate Pure Solvent regeneration R 900 per ton

Sodium hydroxide Diluted in water Regeneration agent R 4 200 per ton

Sulphuric acid 98 mass% H2SO4 Leaching agent R 500 per ton

3.1.4. Processing constraints

It is important to identify and understand the hazardous processing conditions to avoid

equipment damage as well as an unsafe working environment. These process conditions

are known as the processing constraints and should play and integral role in the design of

the plant.

In the leaching section concentrated sulphuric and nitric acid is used which are highly

corrosive. The choice of materials for construction must be resistant to these conditions in

order to improve the economic potential of the project. The dissolution of acids in water is

exothermic and needs to be monitored. The leaching of uranium through a nitric acid

system is autocatalytic and exothermic and this combination may result in hazardous

operating conditions.

An unwanted increase in the pH of the pregnant leach liquor will result in the precipitation of

dissolved impurities and causes fouling of the resin. The resin used in the ion exchange

process is also sensitive the sharp changes in the pH of the system which results in swelling

of the resin. The selectivity of the resin towards uranium is the key aspect in designing ion

exchange equipment.



The organic compounds used in solvent extraction are extremely flammable and cause an

explosion hazard in enclosed vessels. High temperatures will increase the possibility of a

fire or explosion due to the low ignition point of the organic compounds. One of the

important operating conditions is the pH of the system. A high enough pH (above 7) will

result in precipitation of the uranyl sulphate complexes.

The pH of the precipitation section is the main design parameter for the efficient operation of

this section. Small deviations from the optimum pH will result in precipitation of impurities

and product that does not meet specifications.

3.1.5. Other plant and site data

The battery limits and cost of certain facilities is important for the design of a new process.

The availability and capacity of the utilities on a plant needs to be sufficient to supply the

demand. The existing South Uranium Plant already has the required facilities to provide all

the utilities. The utilities used for the entire plant is listed in Table 3.7.

Table 3.7: Utilities

Utility Conditions Price

Steam 175 °C at 12 bar R 50 per ton

Water Demineralised water R 7 per m3

Electricity - 40c per kWh

Compressed air > 4.5 bar -

The waste disposal facilities needed for the tailings from each process unit are tailing dams

and a neutralization plant which complies with international legislation for the protection of

the environment.

3.1.6. Physical properties of all components

The ASPENTech® chemical database is used to create a databank for the physical

properties of all the components used. However not all the components are in the

ASPENTech® chemical database, thus the properties need to be gathered from other

sources which included MSDS data and chemical compound databases on the internet. The

unknown physical properties of the compounds are given in Table 3.8.



Table 3.8: Physical properties of unknown compounds

Compound Specific heat

capacity (J.kg-1.K-1) Vapor

pressure (bar) Molar volume

(cm3/mol)

Heat of formation

(J/mol)

ADU - Solid - 845200 cal/mol

Amine 366 - Stay liquid - -

Brannerite 213 667 Solid 72 -673.119

kcal/mol

Chlorite 561 063.618 Solid 207 -377.081

kcal/mol

Coffinite 136 969 Solid 53 -454.704

kcal/mol

HNO2 - Solid - -87 000

Iron complex resin - Solid 8 000 -

Isodecanol - 0.0013 - -

Kerosene 2 010 0.00689 - -24 149

Loaded Amine 366 - Stay liquid - -

Resin - Solid 7 110 -

U-phosphate 149 105 Solid 76 -465.728

kcal/mol

Uranyl complex resin - Solid 14 489 -

3.2. Level 1: Batch versus continuous

There are three guidelines for the decision between whether a process is batch or

continuous. These guidelines are:

• The production rate of the process per year.

• The market force of the produced material.

• Operational problems that might occur.



To make an informed decision, it is important to consider the influence of all three the

guidelines. The ore capacity for the plant is 4 320 000 ton per year, and according to

Douglas (1988) a capacity greater than 4500 ton per year results in a continuous process.

The product (ADU) is not a seasonal product and is used worldwide which classifies the

process as a continuous process.

The operational problems which occur in the ion exchange process are resolved by

operating the process as semi-batch. Ion exchange is an adsorption and desorption process

and therefore it is operated as a semi-batch process, but this does not affect the continuity of

the entire plant. From this it is concluded that the plant is operated as a continuous process

and the following aspects will help to develop a conceptual design for a continuous process:

• Process units needed

• Interconnections among units

• Estimate the optimum processing conditions

• Which units should be batch or continuous

• Single vessels versus individual vessels for each step

3.2.1. Process units needed

The ore is processed to a slurry by the gold mining section and is the raw material from

which the uranium is extracted. Low grade South African ore is used, with a low uranium

concentration. For this reason the uranium must first be liberated from the ore and then

concentrated to allow easy recovery of the desired product. The four process units required

to recover the desired product are leaching, ion exchange, solvent extraction and

precipitation.

Acid leaching allows the liberation of uranium from the uranium containing minerals in the

ores. Acid is added to the slurry in order to decrease the pH to a desired level at which

uranium is dissolved from the ore. An oxidation agent is also required to oxidize uranium to

the desired oxidation state. The slurry is sent to counter-current decantation where the

liquids and solids are separated. The underflow is sent to the neutralization plant where it is

prepared for the gold extraction process.



The leach liquor has a uranium concentration of 0.15 g U3O8/L which is increased to 4 g

U3O8/L in the ion exchange section. This is achieved by the highly selective adsorption of

uranyl sulphate complexes onto a strong base resin. The uranyl sulphate complexes are

recovered from the resin by elution or desorption with the eluant, diluted nitric acid.

The purpose of the solvent extraction unit is to upgrade the uranium concentration as well

removing impurities. The extraction stage is where the selective dissolving of uranium from

the aqueous phase into the organic phase occurs. Further processing of the organic phase

with water allows the removal of impurities such as iron sulphate complexes. The uranium is

stripped from the organic phase with sulphate ions while the pH is controlled with caustic

soda. The final process unit is precipitation where the solid yellow cake product, ammonium

diuranate, is precipitated. In this unit the specifications of the final product, which include 35

mass% ADU, should be reached.

3.2.2. Interconnections among units

The conceptual design is based on the upgrading of the current South Uranium Plant

situated near Orkney in the North-West Province. The same processing units are used as in

the existing flowsheet and are illustrated as a simplified block flow diagram in Figure 3.1.

Figure 3.1: Simplified block flow diagram of process

The ore is fed to the leaching stage which includes the two counter-current decantation

trains with five thickeners and a clarifier in each train. The leaching slurry is separated into

the pregnant leach liquor and solid slurry which is sent for further processing at the gold

extraction plant. The pregnant leach liquor, which contains less than 50 ppm solids, is sent

to the ion exchange stage.



The concentration of the uranyl complexes in the leach liquor is upgraded and partially

cleansed of impurities in the ion exchange stage. The eluate from the ion exchange is sent

to the solvent extraction unit, where further up concentration and cleansing of the uranyl

complexes takes place. The OK liquor is sent to the precipitation unit where the ADU

product is formed.

3.2.3. Estimate the optimum processing conditions

The optimum process conditions will ensure balance between favourable economic potential

and efficient overall production of the ADU product. These operating conditions are acquired

from the current South Uranium Plant and external literature. The processing conditions for

each stage are given in Table 3.9.

Table 3.9: Processing conditions

Process conditions Value

Leaching

pH (-) 2

Temperature (°C) 30

Final underflow SG (-) 1.55 to 1.6

Ion exchange

pH (-) 1 to 2


Solid content (ppm) < 50

Solvent extraction

pH (-) 2 to 5


Precipitation

pH (-) 7.5


Table 3.9 is the optimum conditions reported from literature, but it is not always possible to

operate the entire plant at these conditions.



3.2.4. Additional conceptual design information

As discussed in the above sections, the overall process is considered to be continuous.

However, some of the processing units are operated as semi-batch units. This section

discusses additional process information regarding the continuity and vessel configuration of

the processing units.

Leaching is a continuous process which requires numerous pachuca tanks in series to

ensure adequate residence time. The vessels in series allow for better mean residence time

and mixing of the entire stream, and when combined ensure optimum conversion.

Manufacturing and designing a single leaching vessel with the capabilities to process the

entire stream is not economically viable. Therefore tanks in series are the optimum

configuration. Counter-current decantation is a solid-liquid separation unit which is usually

done using separation stages. Separate thickeners are used to ensure optimum recovery of

the liquid.

In the ion exchange process the resin remains stationary in fixed-bed columns while the

leach liquor flows through it in the adsorption stage and the eluant flows through it in the

elution stage. Since this is a semi-batch process, only three or four fixed-bed columns are

required. Usually the elution or wash stage is conducted in one column while adsorption

continues in the other columns. Since each column must carry out adsorption, washing and

elution, intensive piping is required.

Solvent extraction is a liquid-liquid separation unit which is operated continuously. The

liquid-liquid separation requires different sections which include extraction, scrubbing,

stripping, and regeneration. Each stage requires a certain amount of stages depending on

the through-put of the system. Precipitation is a fairly simple process which only requires a

vessel that provides enough residence time for adequate precipitation and seeding. It is

possible to uses vessels in series to increase the residence time. The precipitated slurry is

sent to a thickener from which the overflow is sent back to solvent extraction and the

underflow is sent to a centrifuge. The centrifuge continuously delivers a product with a high

concentration ADU by removing liquids.



3.3. Level 2: Input-output structure of the flowsheet

It is important to account for all the chemical compounds in the system to ensure that there

is no build-up or unnecessary reagent and product losses. To ensure accountability, an

input-output structure for the entire process is done. In Douglas (1988) Level 2, the following

questions are used as a guideline to develop the block flow diagram (Douglas, 1988: 118).

• Should the feed stream be purified?

• Should a reversible by-product be removed or recycled?

• Should a gas recycle and purge stream be used?

• Should the reactants be recovered and recycled?

• How many product streams will there be?

• What are the design variables for the input-output structure, and what economic

trade-offs are associated with these variables?

These questions are answered in the following sections.

3.3.1. Feed purification

The major feed impurity exist in the ore feed and the impurities include pyrite, chlorite,

quarts, muscovite, etc. The above impurities are not inert and consume a large amount of

raw materials. It is impossible to economical remove these impurities from the feed stream.

The processing of uranium is of such a nature that the impurities are effectively and

economically removed throughout the process.

All the other raw material feed streams do not include significant amounts of impurities.

Most of the raw material feed streams are diluted with potable water which is considered as

an inert, but serves a purpose in the processing of uranium. Therefore there is no

purification of the feed.



3.3.2. Recycle streams

In Douglas (1988) it is stated that 99% or more of the valuable materials should be

recovered which is achieved by recycle streams or by-product removal. The recycling of

valuable reactants will result in a decrease of raw material cost and reduce the amount of

chemicals introduces into the environment.

The important recycles are nitrate and sulphate ions, solvents and wash solution. Nitrate

ions are the oxidizing agent in the leaching process while the sulphate ions are used as

lixiviant in the leaching process and a complexing agent for the uranyl ions.

It is very important to recycle more than 99% of the organic solvent due to the negative

economic impact. The wash solution is used as a replacement for potable water in the

counter-current decantation section which will recycle un-reacted uranium leading to an

increase in the overall recovery of uranium.

3.3.3. Removal and purge streams

Removal and purge streams are necessary to prevent build-up of impurities in the system.

The three important waste by-products are nitrogen oxide gas, iron sulphate and silicon

dioxide. The full removal of these compounds will prevent a build-up in the system. It is

important to remove these substances in an environmentally friendly and economically

viable manner.

Nitrogen oxide produced in the leaching process is classified as a light component according

to Douglas (1988) which should be purged from the system. Iron sulphate is present in the

aqueous phase and if not removed will affect the product purity. Silicon dioxide is a fouling

agent for the resin used in the ion exchange section which builds-up if not removed.

3.3.4. Number of product streams

It is important to keep track of all the components in the system to prevent impurity build-up

and this is done with an overall input-output structure. The overall input-output structure for

the process is given in Figure 3.2.



Figure 3.2: Input-output structure of overall process

Due to the numerous amounts of components in the system, it is irrelevant to list all the

species in the product and waste streams. Table 3.10 shows the product streams from

Figure 3.2.



Table 3.10: Product stream classification

Stream Destination

ADU Primary product

Slurry to gold plant Valuable by-product

Nitrogen oxide Purge

Iron sulphide Waste

Silicon dioxide Waste

Nitrate ions Recycle

Sulphate ions Recycle

Solvent Recycle

Wash solution Recycle

The ADU and the slurry to the gold plant streams are the only streams that hold economical

value. The final product in the ADU stream is sent to NUFCOR for further processing while

the gold is extracted from the slurry at the gold plant. The waste is sent to the waste

treatment facilities.

3.3.5. Preliminary economic potential analysis

According to Douglas (1988) the economic potential 1 (EP-1) and EP-2 should be done to

ensure that the process is economically viable to proceed to the next design steps. The

EP’s are based on a preliminary mass balance with the following assumptions made:

• H2SO4 is used to control pH in leach section.

• Enough HNO3 is used to oxidize minerals.

• Reactions occur in series.

• Loss of 1.1% of solvent per hour.

• Overall product recovery of 76%.

The preliminary mass balance is based on the input-output structure displayed in Figure 3.2.

Table 3.11 consist of the summarised mass balance which includes reagent and product

prices.



Table 3.11: Level 2 preliminary mass balance with prices

Component Amount (ton/annum) Price per unit (R/ton) Total cost (R/annum)

Raw materials

Alamine (Alkyl amine) 0.8 57 260 45 808

Ammonia 140 15 000 2 100 000

Caustic soda 18 4 200 75 600

Isodecanol 0.3 22 410 6 723

Kerosene 410 m3 5 000 per m3 2 050 000

Magnafloc 90L 86 22 500 1 935 000

Nitric acid 19 000 1 500 28 500 000

Ore 4 320 000 - -

Potable water 8 600 m3 7 per m3 60 200

Resin (Ambersep TM 400) 62 57 000 3 534 000

Sulphuric acid 54 300 500 27 150 000

Products

Uranium 1068 743 786.41 794 363 885.88

Gold 2.16 254 552 770 549 833 983.2

The EP-1 analysis is the income from the product minus the expense of the raw materials

used per annum. The EP-1 analysis gives a positive value of R 728 million per annum. The

EP-2 analysis is the income from the product and by-product minus the expense of the raw

materials. It is assumed that 0.5 g gold per ton of ore feed is produced. The EP-2 analysis

gives a positive value of R 1 278 million per annum. The EP-1 and EP-2 both gives large

positive values, thus the design process can continue to Douglas level 3.

3.4. Level 3: Recycle structure of the flowsheet

The degree of recycling plays an important role in the economical potential, sizing and

operational cost of equipment. In Douglas (1988) it is stated that a number of decisions

need to be made concerning the recycle structure of the flowsheet. These decisions are

described in the following sections.



3.4.1. Reactor systems required

The process of extracting uranium consists of four sections, with an additional two sections

for the counter-current decantation after the leaching and precipitation sections. In this

section, the leaching and precipitation is regarded as the only reactor systems. The ion

exchange and solvent extraction sections have a high selectivity and recovery of the desired

product complexes and for this reason are regarded as efficient separation systems. Figure

3.3 shows the reactors as white blocks and separation systems as black.

Figure 3.3: Block flow diagram for reactor system

In Figure 3.3 the white blocks represents the reactor systems while the black blocks

represents the separations system as discussed above. The product streams of the various

systems contain valuable reagents which can be recycled to reduce raw material costs.

3.4.2. Number of recycle streams

The valuable reagents which leave the process include streams with high levels of nitrate,

sulphate, and ammonium ions. These streams are recycled to specific process units where

it will most effectively reduce the raw material cost. These recycle structures are shown in

Figure 3.4.



Figure 3.4: Recycle structure

In Figure 3.4 three recycle streams are indicated which hold economical and environmental

advantages for the process. The first recycle stream is the barren liquor form the ion

exchange section which is stripped of uranyl sulphate complexes. The stream contains

large amounts of water and is therefore recycled to the counter-current decantation section

as wash solution. Make-up wash water is still required, but in much less quantities to

effectively remove the uranyl sulphate complexes from the remaining solid slurry. The

reduction in make-up wash water results in less potable water being contaminated.

The second recycle stream is the barren eluate from the ion exchange section. At the ion

exchange section sulphate ions in the eluant are used for desorption of uranyl sulphate

complexes from the resin. An excess of sulphate ions are fed to ensure complete removable

of the uranyl sulphate complexes due to the equilibrium of the system. The excess sulphate

ions are removed in the extraction section of solvent extraction and are recycled back to the

leaching process which will reduce the amount of lixiviant needed.

The third recycle stream is the aqueous solution from the solid-liquid separation stage at the

precipitation section. This stream contains ammonium and sulphate ions which is required

as raw materials for the precipitation and solvent extraction respectively. These three

recycle streams play an important role in the economical potential of the process.



3.4.3. Abundance of reactants

An abundance of reactants may cause virtually full conversion of the desired reactant, but

will have a negative impact on the economic potential analysis. There are numerous side

reactions for the hydronium and nitrate ions which forms unwanted by-products. An

abundance of the required reactant is supplied to the leaching section to ensure a maximum

recovery of the uranium while sustaining the unwanted side reactions.

The ore feed consist of many impurities which cannot be controlled or separated. These

impurities consume the reactants and form unwanted products. Hydronium ions are used to

dissolve most of the ore minerals while nitrate ions acts as an oxidising agent for the

uranium containing minerals and pyrite. Since the dissolution of uranium bearing minerals

are the main objective, the oxidising agent is fed in excess. The hydronium ions are

controlled at a certain level to ensure the desired pH in the leaching system.

The hydronium ions are mainly supplied by the addition of sulphuric acid which is more

economical compared to nitric acid. This results in an excess of sulphate ions in the

leaching section which is required for the adequate formation of uranyl sulphate complexes.

3.4.4. Operational considerations

Specific utilities for reactors are cooling water, steam, electricity and pressurised air.

Although pressurised air and steam are required at the reactors in the uranium extraction

process. Pressurised air and steam are available from the gold extraction plant which is

located near the uranium plant and therefore there is no need for compressors or boilers on

the South Uranium Plant. The ammonium gas required for the precipitation section is

assumed to be supplied under sufficient pressure. The pressure gradient is used to

transport the ammonium gas to the precipitation section.

In most cases the leaching and precipitation reactors are heated adiabatically to increase the

reaction rates. The immense quantities of water, which is a heat carrier, absorb the small

amounts of energy that is released from the reactions. The temperature difference between

night and day time as well as seasons should be taken into consideration for the operation of

the reactors.



3.4.5. Recycle economic evaluation The economic potential of the project should be continuously evaluated to ensure that the

project does not reach a negative value. The third level of evaluation, EP-3, is done by

subtracting the reactors costs from the EP-2 value calculated (Douglas, 1988: 158). The

reactors in this process include the leaching and precipitation tanks. The total costs of the

reactors should consist of the capital cost for the equipment and the cost of the power

needed to operate the reactors. The operations needed for these reactors are air agitation

and steam heating.

As mentioned this is an upgrade project and therefore existing equipment is used for both

leaching and precipitation therefore the only capital cost required is the purchasing of pumps

and storage tanks for the nitric acid used in the leaching process. Steam is necessary to

heat the pachuca tanks in the leaching unit and the feed stream to the precipitation unit to a

temperature of 30 °C. The steam costs are calculated using a utility cost of R 50/ton and the

results are shown in Table 3.12 together with the capital costs of the equipment.

Table 3.12: Capital and operation costs for reactors

Unit Capital cost (R) Operation cost (R) Total (R)

Leaching 34 681 483.46 10 306 799.20 44 988 282.66

Precipitation 15 770 282.59 10 005 643.27 25 775 925.86

Total: 70 764 208.52

Using the capital and steam costs calculated in Table 3.12, the EP-3 is determined to be

R 1 207 million, which is an indication of a positive economic potential for this stage of the

feasibility study. This economic evaluation shows positive results, therefore the study can

continue to the next phase.



3.5. Level 4: Separation systems

The usual Douglas (1988) design hierarchy is followed for Levels 0 through 3 for the mineral

extraction process. The usual Douglas (1988) design hierarchy is mainly used for

petrochemical plants; however the main difference in mineral extraction is the presence of

solids in the system. For this reason the Rossiter and Douglas (1988: 408) Level 4 is

followed for the solids system. The questions that need to be answered for this level are

(Douglas, 1988: 410):

• How can the primary product be recovered?

• What type of solids recovery systems is used?

• How should the waste-solid separation be accomplished?

• Are any liquid-liquid separations required?

• Locations of separation units (purge or recycle streams or both)?

For mineral extraction processes three physical phases are present, and each phase has its

own recovery system. Level 4 can by divided into the following sections:

• General structure.

• Vapor recovery system.

• Solid recovery system.

• Liquid recovery system.

3.5.1. General structure

The first separation unit occurs in the leaching section, where two products streams leave

the system. One stream is the solid-liquid stream which contains the desired product and

by-product. The second product steam is a result of the air agitation used and may contain

nitrogen oxide gas. The amount of nitrogen oxide formed is completely soluble in the

aqueous phase. The solid-liquid product stream is processed in separate separation units

and a general separation structure is given in Figure 3.5.



Figure 3.5: General separation structure

The solid-liquid product leaving the leaching section proceeds to the counter-current

decantation section where the liquids are recovered from the leaching product stream. The

liquid stream from the counter-current decantation section is sent to the ion exchange

system where the uranyl sulphate ion complexes are selectively removed to the eluate. The

eluate then flows to the solvent extraction unit where impurities are removed using liquid-

liquid separation. The stripping liquor from the solvent extraction section is sent to the

precipitation section. The slurry from the precipitation section is sent to a solid-liquid

separation unit where the ADU product is recovered.

3.5.2. Vapour recovery system

Vapour recovery systems are used to recover valuable gasses from a purge or vent stream.

No vapour recovery system is necessary because small amounts of valuable gasses are

present in the vent. It is not economically viable to recover these small amounts of the

valuable gasses. The vent stream will mainly consist of air due to air agitation.

3.5.3. Solid recovery system

From Figure 3.5 it is noted that there is two solid recovery systems in the uranium extraction

plant. The first is the waste solid recovery which is used to remove the un-reacted ore

containing the valuable gold by-product. The second solid recovery is where the ADU

product is separated from the aqueous phase.

Counter-current decantation is used in the waste solid recovery which is done in two

identical thickener trains. The waste solid recovery system is shown in Figure 3.6.



Figure 3.6: Waste solid recovery system

The first thickeners in each train serves as a clarifier to ensure low solid content in the

overflow sent to ion exchange. The floculant, Magnafloc 90L, is distributed evenly over the

five stages of the solid recovery system for the gold, but is not added to the first thickener.

This counter-current decantation system ensures a uranium recovery of 99.99%. The slurry

sent to the gold extraction plant is washed with the barren leach liquor from the ion

exchange section.

The second solid recovery system is the ADU precipitation unit and is illustrated in Figure

3.7. The OK liquor from solvent extraction is sent to the two tanks in series where

precipitation of the ADU takes place. The solution containing the precipitated product is sent

to a thickener to settle the ADU cake and decant the ammonium sulphate which is sent back

to the solvent extraction stripping section. The thickener underflow, containing ADU, is sent

to two centrifuge stages. Inside the first centrifuge stage the slurry is sent to two centrifuges

in parallel where the ADU is washed with demineralised water spray. The solids are sent to

the second stage centrifuge while the liquids are pumped to the centrate tank so that it can

be recycled to the ADU thickener. The solids from the second centrifuge stage flows to the

final product storage tanks for transport to NUFCOR while the liquid is sent to the waste

water treatment facility.



Figure 3.7: Solid recovery system for ADU

3.5.4. Liquid recovery system

Figure 3.5 show that there are two liquid recovery systems in the uranium extraction process

with the main purpose to purify and increase the concentration of the U3O8. The increase in

concentration and purification allows more economical recovery of the desired product. Ion

exchange is the first process used for the purification and up-concentration of the product,

and the second is the solvent extraction process.

The leach liquor contains low concentrations of uranyl sulphate complexes with a high

volumetric flow. The ion exchange process is a liquid recovery system where the uranyl

sulphate ions are transferred from the leach liquor to the eluate which has a much lower

volumetric flow and a higher uranyl sulphate concentration. This is achieved by adsorption

of these ions onto resin particles and then eluted from these particles into the eluate. The

adsorption and elution of the resin takes place in the same columns at different stages. This

concept is shown in Figure 3.8.



Figure 3.8: Schematic representation of ion exchange

Figure 3.8 is a schematic representation of the operation of fixed-bed column ion exchange

where column 1 and 3 is in the adsorption stage while column 2 is being eluted. Before

each elution stage, the resin is washed with water to remove solid particles. This setup

usually contains a total of three or four columns.

The second liquid recovery system is the solvent extraction section. The solvent extraction

process requires four different sections to ensure efficient separation between the aqueous

feed and the solvent added. These sections are extraction, scrubbing, stripping and

regeneration; the sections are displayed in Figure 3.9.

Figure 3.9: Schematic representation of solvent extraction



The extraction section is responsible for the extraction of the uranyl sulphate complex from

the aqueous phase into the organic phase which will include entrainment of iron sulphate

complex. The scrubbing section removes the entrained iron sulphate complex by washing

the solution with demineralized water. The last section is where the uranyl sulphate complex

is stripped from the organic solvent with ammonia sulphate and up-concentrated in the

aqueous phase. The solvent is regenerated in one stage with caustic soda and sodium

carbonate to remove any impurity build-up. Each section contains multiple stages to

increase the recovery of U3O8.

3.5.5. Separation system economic evaluation

The final economic evaluation level in the design procedure is level four (EP-4) in which

purge losses and separation unit costs are considered (Douglas, 1988:189). In the designed

process no products are purged, only reagents. The cost of these reagent losses is

considered in the raw materials cost estimations thus no purge loss calculations is

necessary. The separation systems needed for the process, as mentioned above, is

counter-current decantation, ion exchange, solvent extraction and the ADU solid separation

system. The capital and operation cost for the equipment and pumps of the separation

systems are calculated and given in Table 3.13.

Table 3.13: Separation systems cost calculations

Unit Capital cost (R) Operation cost (R) Amount (R)

CCD 69 999 662.07 4 003 586 74 003 248.07

Ion exchange 6 504 779.59 10 294 450 16 799 229.59

Solvent extraction 26 134 497.60 200 215.98 26 334 713.58

Total: 117 137 191.24

The final EP-4 is calculated by subtracting the total capital and operating costs of the

separation units from the EP-3 calculated value. The EP-4 value obtained is R 1 089 million,

which gives a positive economic potential indication. This is the final economic evaluation

which is a good indication of the economic feasibility and therefore the study can continue to

the next phase.



3.6. Level 5: Heat integration

Energy conservation is an important aspect that is considered in the design of a process and

can be done by means of energy integration. During energy integration, heat is transferred

from a hot stream that needs to be cooled, to a cold stream that needs to be heated. The

heat integration is done by analysing the minimum heating and cooling requirements for the

heat-exchanger network (Douglas, 1988:216).

Investigating the current uranium extraction process, energy is only necessary for the

heating of the leaching pachuca tanks as well as the OK liquor stream to 30 °C. The heating

of the leaching feed streams is done with steam and no cooling is needed. Thus heat

integration will be redundant because there are only two heating requirements. Therefore

only one heat exchanger is used at the precipitation unit and no heat integration network is

created.

3.7. Equipment design

In an expansion project it is vital to determine the capacity needed after expansion, which

will allow the evaluation of current equipment re-sizing. The sizing methods followed to size

each of the discussed equipment below is fully described in Appendix B. Since the detail

design of the solvent extraction unit is done in Chapter 4, the solvent extraction equipment is

not designed in this section.

The first process unit, leaching, is sized using reaction kinetics and sizing methods

discussed by Minerals Council of Australia (2006). From this it is calculated that a

conversion of 93% is achieved, utilising 11 out of the 14 existing air agitated pachuca tanks.

The operating temperature for the leach process is optimised at 30 °C and a nitrate

concentration of 0.0618 mol/L , which allows for optimum energy and reagent consumption.

The counter-current decantation section is sized according to the existing equipment. The

unit area is calculated for the existing equipment and falls in the range given by Merritt

(1971). Using the unit area, a second train of 6 thickeners is sized to process the extra feed

capacity. The first thickener in the new train has a diameter of 50 m and the rest have a

diameter of 45 m.



Fixed-bed ion exchange columns are preferred to the existing continuous counter-current ion

exchange process. The capacity of the existing columns is determined using basic mass

balance combined with resin specifications for Ambersep TM400. It is found that the existing

columns have adequate capacity to process the increased feed. Four of the existing

adsorption columns are used while the remaining two are back-up columns. The existing

elution column is used for regeneration of the resin.

The precipitation reactor is sized using kinetics from Table 3.4. The two existing

precipitation reactors are adequate to achieve high efficiency thus it is assumed that the

tanks and centrifuges is large enough to handle the increased capacity. The precipitation

thickener is sized using the same method as for the counter-current decantation. The

existing reactors, thickener, tanks and centrifuges have sufficient capacity to process the

increased feed. Table 3.14 gives a summary of the numbers and sizes of the equipment

used for the plant.

Table 3.14.a: Summary of leaching equipment

Specifications Description

Construction material Rubber lined stainless steel

Diameter per tank 10 m

Feed rate 554

Number of tanks 11

Residence time per tank 1 hour and 21 minutes

Tank alignment Series

Volume per tank 750 m3



Table 3.14.b: Summary of counter-current decantation equipment


Construction material Rubber lined cement

Trains 2

Diameter per tank Thickener 1a : 60 m

Thickener 2a – 6a: 55 m

Thickener 1b: 50 m

Thickener 2b – 6b: 45 m * With subscripts a and b being train 1 and train

2 respectively

Feed rate Train 1: 340.7 m3/hr

Train 2: 212.9 m3/hr

Number of stages 6 per train

Residence time per stage Thickener 1a : 33 hr and 10 min

Thickener 2a – 6a: 28 hr

Thickener 1b: 35 hr and 20 min

Thickener 2b – 6b: 28 min and 30 min * With subscripts a and b being train 1 and train

2 respectively Volume per tank Thickener 1a : 11 300 m3

Thickener 2a – 6a: 9 500 m3

Thickener 1b: 7 500 m3

Thickener 2b – 6b: 6 000 m3 * With subscripts a and b being train 1 and train

2 respectively



Table 3.14.c: Summary of ion exchange equipment


Adsorption time 20 hr

Construction material Glass reinforced plastic

Diameter per tank Adsorption columns: 3 m

Regeneration columns: 3 m

Elution bed volumes 7

Elution time 7 hr and 45 min

Feed rate 570 m3/hr

Number of tanks Adsorption columns: 6

Regeneration columns: 6

Residence time per tank Adsorption: 2 min and 14 sec

Elution: 20 min




Table 3.14.d: Summary of precipitation equipment


Reactor

Construction material Stainless steel 316L

Diameter per tank 1.5 m

Feed rate 7.8 m3/nr

Number of tanks 2

Residence time per tank 12 min and 20 sec

Volume per tank 2.667 m3

Thickener

Construction material Stainless steel 316L

Diameter per tank 15 m

Feed rate 7.8 m3/hr

Number of stages 1

Residence time per stage 78 hr


3.8. Mass and energy balance

The mass and energy balance over the entire plant is required in order to validate the

technical feasibility of the process. The mass and energy balance are in some cases

interdependent, since the energy flow is calculated from the mass flow and the mass

conversion rate are dependent on the energy of the system.



3.8.1. Mass balance

The mass balance for this process is done to account for all the materials in the process.

The law of conservation states that mass can not be created nor destroyed. An overall mass

balance is given in this section while more detailed mass balances are given in Appendix A.

Table 3.15 is the result obtained from the overall mass balance. This overall mass balance

does not consider internal recycles for the resin and solvent.

Table 3.15: Overall mass balance

IN OUT

Stream kg/hr Stream kg/hr

Ore feed 880 000 Solid slurry to gold plant 992 487.8

HNO3 feed 2 497.44 Removed silicone from resin 63.781

H2SO4 feed 15 320.9 Spent demin. water 5 610

Wash water make-up 0.998 Precipitation recycle bleed 401.49

Slaked lime 67 584.13 Spend liquid from centrifuges 1 297.82

Eluant 27 507.32 ADU product 374.012

Regeneration feed 25.352

Demin. water 5 600

Centrifuge wash water 1 410.02

Ammonia 85.73

Total 1 000 036 Total 1 000 239

The total amount of materials that entered the process adds up to 1 000 036 kg/hr while the

materials leaving the process is 1 000 239 kg/hr. Comparing these two values, an error of

0.02% is found. This error occurs due to the simulation program used to simulate the

extraction process, ASPENTech®, which uses tolerances to simulate the mass balance and

therefore this error is acceptable.

3.8.2. Energy balance

The energy balance gives an indication of the thermodynamical feasibility of the process

indicating energy requirements, energy generation and energy losses with detailed energy

balances supplied in Appendix A. In this section the energy balance is done with the

following assumptions:



• Heat capacity of solids: 0.82 kJ/kg/K

• Kinetic energy is negligible

• Heat of reaction and heat of dissolution are negligible

There are only two processes on the plant that will require thermal energy; the leaching

pachucas and the precipitation reactors. Energy loss to the environment generally occurs at

the leaching pachucas, and thickeners at leaching and precipitation. The reactions in all the

processes take place in small amounts due to diluted mixtures and therefore are not taken

into account for energy generation or consumption. However the reactions in the

neutralisation process may have a significant effect on the temperature of the mixture (see

Appendix A).

The energy required for this plant operations are 10.8 MW at the leaching section and 0.75

MW at the precipitation section. From calculations it is found that the energy generation

caused by reactions in the neutralisation and precipitation processes has no significant effect

on the temperature of the product streams. It is concluded that it is impractical to implement

heat integration systems on this plant.

3.9. Process Flow Diagrams and process description To conclude the process design procedure, a detailed process desctription is provided. The

project entails the expansion of the South Uranium Plant (SUP) of AngloGold Ashanti to

process an increased ore feed from 240 000 to 360 000 ton per month. The operating time

for this plant is planned at 8 150 hours per year and the production rate for ADU is

calculated to be 1 068 ton per annum. This upgrade of the SUP is done by evaluating the

existing equipment and if necessary, design new equipment with the capacity to process the

new feed. The following process units are defined and illustrated in detail on the process

flow diagrams (PFD’s) and discussed below.

• Unit 1 (U01): Leaching, counter-current decantation and neutralisation

• Unit 2 (U02): Ion exchange

• Unit 3 (U03): Solvent extraction

• Unit 4 (U04): Precipitation



PFD unit 1: Leaching and CCD



PFD unit 2: Ion exchange



PFD unit 3: Solvent extraction



PFD unit 4: Precipitation



3.9.1. Unit 1 (U01): Leaching, CCD, and neutralisation

The initial stage in the extraction of uranium is the leaching stage followed by counter-

current decantation (CCD) of the leach liquor. The ore feed is received from the crushing

and milling circuit with an added 330 000 kg/hr of water to give the ore feed a SG of 1.6 and

total mass flow of 880 000 kg/hr. The ore is fed to the first leaching tank in the series to

extract the uranium from the ore. The leach product is sent to the CCD to recover the

uranium containing liquids and to settle out the solids containing gold which is sent back to

the gold plant after neutralization. The liquid overflow from CCD contains the valuable

uranium and is sent to the ion exchange unit for further processing. Unit 1 can be divided

into three sections which are discussed below.

• Leaching

• CCD

• Neutralization

Leaching (U01-P01 to U01-P13)

The leaching process is done in a series of 11 air-agitated and steam heated leaching

pachucas (U01-P01 to U01-P11) where an oxidant and lixiviant is added to liberate the

uranium from the ore. The air agitation is done to ensure efficient mixing of the solution in

the pachucas and the steam is added to every fourth pachuca. The temperature of the first

pachuca (U01-P01) is raised to 30°C, using steam addition. It is very important to ensure

that the temperature throughout the leaching section is controlled correctly to ensure

optimum leaching kinetics. The existing pachucas is 750 m3 which provide an adequate

capacity to process the upgraded feed due to the change in oxidant used. A residence time

of 1.5 hours is achieved for each pachuca tank. The uranium conversion in the leaching

section is 96% for the mineral uraninite. This conversion is reached due to the faster

leaching kinetics of the nitric acid leaching system.

The oxidant used is nitric acid which oxidises the uranium ions from U4+ to U6+. U4+ is

insoluble in aqueous systems but U6+ is soluble. An added benefit of the nitric acid is that

multiple reactions occur which create an auto-catalytic environment resulting in faster

kinetics and less oxidant needed. Sulphuric acid is used as lixiviant to control the pH of the

leach product at 2. Sulphuric acid is used to control the pH, rather than nitric acid, because

two hydronium ions are formed per molecule of sulphuric acid which is reacted. The ore is



fed to either U01-P01 or U01-P02 using a splitter box, depending on the operatability of

UO1-P01.

It is important to control the nitrate ion concentration and temperature of the solutions in the

pachucas to ensure optimal process conditions and efficient leaching. The reagent flow to

the puchacas should be controlled by the amount of ore that is received to prevent abundant

used of expensive reagents when the ore feed. It is suggested that the level of agitation is

regulated to ensure efficient mixing of the solution in the pachucas to ensure adequate

leaching efficiency.

CCD (U01-TH01A to U01-TH06A & U01-TH01B to U01-TH06B)

After the liberation of the uranium in U01-P13 the leach product flows to the thickener

capacitance tank (U01-DM01). U01-DM01 has a volume of 750 m3 and is also air agitated

to ensure that no settling occurs in this vessel. From U01-DM01 two streams are pumped to

separate thickener trains, Train A (U01-TH01A to U01-TH06A) and Train B (U01-TH01B to

U01-TH06B). Raked gravity thickeners are used to separate the solids, which are sent to

the gold plant, from the liquids which are sent for further processing at ion exchange. In

each train the first thickener (U01-TH01A and U01-TH01B) serves as a clarifier to ensure no

solids is sent to the ion exchange. The clarifiers for Train A and B are 60 and 55 meters in

diameter respectively while the rest of the thickeners in the trains are 50 and 45 meters

respectively.

The leach product (U01-S17) that is fed to Train A is 62 wt% of the entire stream while Train

B is 38 wt%. The thickeners are operated in a counter-current configuration as displayed in

Page 1 of 4 of the PFD’s. The flocuant (Magnafloc 90L) is added to the thickeners to

promote faster settling of the solids to achieve effective separation. The floculant is

distributed between the five thickeners in each train, but no floculant is added to the

clarifying thickeners. The wash solution added to the thickeners is recycled from the

adsorption columns at ion exchange together with a potable water make-up stream. A wash

ratio of 1:1 is used to ensure that less than 0.01 wt% of the uranium is lost. The overflow of

U01-TH01A and U01-TH01B, known as the pregnant leach liquor, is sent to the ion

exchange unit (U02), while the underflow of U01-TH06A and U01-TH06B, containing the

gold, is pumped to the final pachuca (U01-P14) for neutralization of the slurry..

The efficiency of the thickeners depends on the residence time and amount of wash solution

added to the thickeners. The solid content of the pregnant leach liquor should not be more



than the prescribed 50 ppm as the solids can poison the resin used in ion exchange. It is

suggested that the height of the rake inside the thickeners should also be controlled

because the thickness of the slurry might damage the equipment.

Neutralization (U01-P14)

The neutralization facility is needed to ensure that the pH of the slurry, which is sent to the

gold plant, is maintained at 10.5 by the addition of slaked lime. The neutralization is done in

an air-agitated pachuca to ensure efficient mixing and neutralization of the slurry. Not all of

the existing pachucas are used for the leaching process therefore the last existing pachuca

(U01-P14) is used for neutralization.

3.9.2. Unit 2 (U02): Ion exchange

The liquid product from U01 contains the valuable uranium ions and other impurities. The

ion exchange process is used to remove these impurities and increase the concentration of

uranium ions from 0.2 to 4 g U3O8 per litre. In this process uranyl sulphate ions are

adsorbed onto resin, the loaded resin is washed to remove impurities and finally the uranyl

sulphate ions are stripped from the resin during elution. All of these processes, except for

the regeneration stage, are done in the same column, leaving the resin stationary inside the

column. The resin attrition in this system is reduced, since the resin is not moved frequently.

This unit can be divided into four steps listed below.

• Adsorption

• Wash

• Elution

• Regeneration

Adsorption (U02-AC01 to U02-AC05)

The pregnant liquor stream (U02-PS01) entering U02 is sent to the first adsorption column

(U02-AC01) from where it is pumped through the rest of the columns (U02-AC02 to U02-

AC05) in the adsorption train. Each of these columns has a capacity of 145 m3. Due to the

selectivity’s of the sulphate ions and uranyl- and iron sulphate complexes, these complexes

are mainly absorbed onto the resin (Ambersep TM400).



The selectivity’s allow for 60% of the resin capacity to be occupied by the uranyl sulphate

complexes, 10% is occupied by iron sulphate complexes and the remaining 30% by

sulphates. The adsorption front as discussed in Section 2.5.2 moves downward through the

leading adsorption column. When uranyl sulphate complex content of the exit stream of a

adsorption column is equal to that of the feed stream, the resin in the column is saturated

with uranium and the column is isolated from the adsorption train. The stream exiting the

adsorption column train is the barren liquor, which is recycled back to CCD as wash solution.

The use of the fixed-bed ion exchange process means that smaller columns is necessary to

achieve the desired adsorption, therefore five of the existing ion exchange columns will be

used. These columns will be modified by covering the columns with a lid to allow the back-

wash process to be done in these columns. It is calculated that each adsorption will initially

take approximately 30 hours to reach saturation.

Wash (U02-AC01 to U02-AC05)

Once the column has been isolated, the wash process is started. In the wash process any

impurities and unwanted solids are removed by fluidising the resin with potable water which

is fed from the bottom of the column. Care must be taken to avoid blowing out the resin with

the wash water. The wash water is recycled back to the pregnant leach liquor feed tank to

recover any pregnant leach liquor and therefore uranium from the saturated resin.

Elution (U02-AC01 to U02-AC05)

The elution process commences after the wash process is completed. This process is

necessary for desorption of the uranyl sulphate complexes from the resin into the aqueous

phase. The eluant used for desorption is sulphuric acid which is diluted in U02-DM01. The

eluant is pumped from the top of the column, exiting as the uranium rich eluate. This eluate

is sent to solvent extraction for further processing. The eluate flow required is determined by

the volume required to reach the desired uranyl sulphate concentration. It is estimated that

eleven bed volumes of eluate is used to elute each resin column with a residence time of

approximately 15 minutes.



Regeneration (U02-RC01)

The leach liquor contains certain ions that forms a strong bond to the resin and are not

removed during the wash stage. These ions are referred to as poisons but can be removed

by regenerating the resin as discussed in Section 2.5.2. Regeneration is done once every

two days by removing the one column from the train, draining the resin out of the column

and pumping the resin to the regeneration column (U02-RC01).

The regeneration step is complex and consists of several internal steps to protect the resin

from the damage caused by pH-shock. The resin from the acidic system is first contacted

with potable water, then with diluted caustic solution and finally caustic solution. This is

done to ensure a gradual pH increase and prevent pH-shock. Before the resin is returned to

the adsorption columns, the pH is gradually decreased by reversing the above procedure.

Some permanent poisons are not removed during regeneration and the resin capacity

decreases over long periods. For this reason the resin inventory is replace with new resin

every three years as suggested by Merritt (1971).

3.9.3. Unit 3 (U03): Solvent extraction

The solvent extraction unit (U03) is included in the extraction process to increase the

concentration of uranyl sulphate ions and remove the iron sulphate complexes which allows

for more efficient precipitation. Solvent extraction is a form of liquid-liquid separation in

which an organic and aqueous phase is used. There are four steps in the solvent extraction

process which are each discussed below.

• Extraction

• Scrubbing

• Stripping

• Regeneration



Extraction (U03-MSE01 to U03-MSE03 & U03-AS01)

The eluate received from U02 is stored in a 600 m3 tank (U03-ST01), from where it can be

continuously fed to the first mixer-settler (U03-MSE01) in the extraction section. The

aqueous phase is pumped to the mixer from below from where the impeller pumps the

solution further. The mixer-settlers are operated in a counter-current configuration with the

organic phase fed to the last mixer-settler (U03-MSE03) in a ratio of 1.1:1 (Vorg:Vaq). The

after-settler (U03-AS01) is used to reduce solvent loss, thus no organic is fed to the mixer-

settlers.

The uranyl sulphate complexes are extracted from the aqueous phase into the organic

phase. In addition to the uranyl sulphate extraction, iron sulphate complexes are entrained

in the solvent as an impurity. The raffinate (barren aqueous solution) from U03-AS01 is

recycled to leaching since it contains a large amount of sulphate and hydronium ions.

Scrubbing (U03-MSS01 to U03-MSS03)

The scrubbing section is necessary to remove any impurities contained in the solvent. The

scrubbing is done with demineralised water in a counter-current configuration. The organic

solvent from U03-MSE03 is fed to U03-MSS01 while the demineralised water is fed to

U03-MSS03 from a demineralised water storage tank (U03-ST04). In this section the ion

sulphate complexes entrained in the solvent, is washed out with demineralised water.

Stripping (U03-MST01 to U03-MST04 & U03-AS01)

The uranyl sulphate complexes are desorbed from the organic phase back to the aqueous

phase in the stripping section using ammonia sulphate which is recycled from the

precipitation unit (U04). If the pH in this section increases, the product will begin to

precipitate therefore the pH must be controlled between 2.5 and 5 by the addition of caustic

soda. The ammonia sulphate entering the stripping section has a high pH which makes it

difficult to control the pH in U03-MST01. Therefore the pH is only controlled in U03-MST02

to U02-MST04 and U03-MST01 is an extra mixer-settler used to increase the efficiency of

the stripping section. The after-settler (U03-AS02) is used to reduce solvent loss, thus no

organic is fed to the mixer-settlers. The aqueous phase, containing the uranyl sulphate

complexes, exiting U03-AS02 is called the OK liquor. The OK-liquor is sent to the

precipitation for the final extraction.



Regeneration (U03-MSR01)

The solvent exiting U03-MST01 proceeds to U03-MSR01 where it is regenerated to ensure

maximum efficiency of the solvent. Sodium carbonate and caustic soda is diluted into a

regeneration solution in U03-MU02 using demineralised water. The regenerated solvent is

recycled back to the organic storage tank (U03-ST03). The regeneration solution is re-used

until it is exhausted. When the regeneration solution is spent, a valve in stream U03-S12 is

opened into a storage tank (U03-ST08) from where it is sent to the waste treatment facility.

3.9.4. Unit 4 (U04): Precipitation

The final unit in the uranium extraction process is the precipitation process unit which is

necessary to form a solid product with the desired specifications. The precipitation of ADU is

achieved by adding ammonia gas to the OK liquor containing uranyl sulphate ions. The final

ADU product is sent to NUFCOR for further processing. This unit can be divided into three

sections which are discussed below.

• Precipitation

• Settling

• Washing and drying

Precipitation (U04-PT01 & U04-PT02) The precipitation of ADU is done in two vessels, the first in which nucleation occurs and the

second in which crystal growth occurs. Crystal growth is important because no floculant is

used for the settling of the solids therefore the ADU particles should be substantially large to

settle naturally. To ensure sufficient formation of particles two vessels are used.

The OK liquor received from solvent extraction is heated to 30°C in a heat exchanger (U04-

HX01) using steam at 175°C and 12 bar. The heated OK liquor flows into the first

precipitation tank (U04-PT01) where ammonia is mixed with compressed air and added to

U04-PT01 from below. From U04-PT01 the entire slurry flows to the second precipitation

tank (U04-PT02) where ammonia and compressed air are added again. The ammonia is

added in combination with the compressed air to ensure even distribution of the ammonia in

the vessel and helps with the agitation of the solution. It is important to provide sufficient

agitation in the precipitation vessels to ensure that settlement do not occur in these vessels.

Therefore the vessels are also mechanically agitated. The vessels that are available on the



existing plant, is large enough to process the increased load and therefore the existing 3.5

m3 tanks will be used.

The reactor temperature should be controlled at 30°C and the pH of the slurry should be

controlled at 7.5 to ensure that the optimal precipitation occurs. The OK liquor feed flow rate

and liquid levels in the precipitation tanks should be controlled to ensure adequate residence

time for the ADU particles to form. It is also necessary to control the ammonia gas and

compressed air entering the reactors to ensure the correct ratio of ammonia to air is fed to

the reactors.

Settling (U04-TH01)

The slurry containing the precipitated ADU particles is sent to the ADU thickener (U04-TH01)

to allow the solids to settle in order to separate the ADU cake from the liquid containing

impurities. As mentioned, no floculant is added to the thickener and only one thickener is

sufficient for the necessary separation. The necessary thickener volume is calculated to

ensure sufficient residence time. The calculated value showed that the existing thickener

with a volume of 710 m3 can handle the increased feed.

The overflow of U04-TH01, containing a high concentration ammonium sulphate, flows into a

storage tank (U04-ST01) with a capacity of 35 m3. From U04-ST01 the liquid is pumped

back to solvent extraction where it is used as stripping agent to strip the uranium from the

organic phase. It is necessary to control liquid level in U04-TH01 to ensure enough

residence time for the solids to settle and to produce an effective overflow. The underflow of

U04-TH01 flows to the first stage mixing tank (U04-MT01) and from U04-MT01 the slurry is

pumped to two centrifuges in parallel. The centrifuges are used to wash and dry the ADU

product.

Washing and drying (U04-CG01 to U04-CG03)

The centrifuges are used to wash off the impurities such as SO42- and to continuously

dewater the ADU slurry. The current capacity of the existing centrifuges is 100 – 150 kg

ADU/hr which is sufficient to process the increased capacity.

The slurry from U04-MT01 is pumped to the first stage centrifuge which consists of two

centrifuges in parallel (U04-CG01 and U04-CG02) and demineralized water is added to both

centrifuges in the form of a spray to wash the solids in the slurry. The wash solution from



the centrifuges flows to a storage tank (U04-ST02) and is eventually pumped back to U04-

TH01 to ensure that no uranium is lost in the wash solution. A part of this recycle stream is

bled to the waste water treatment facility to prevent built-up in the system. The solids from

U04-CG01 and U04-CG02 are pumped to the second stage mixing tank (U04-MT02) and

from there it is pumped to the second stage centrifuge U04-CG03. In the second stage

centrifuge no water is added and this centrifuge is only used for the dewatering of the

product. The liquid from U04-CG03 continues to a storage tank (U04-ST03) and is pumped

to the waste water treatment facility. The solids coming from U04-CG03 is pumped to the

final storage tank (U04-ST04) where the final product is continuously stirred and air agitated

to ensure that settlement does not occur in the storage tank.

The flow of the ADU slurry to all the centrifuges should be controlled to ensure optimal

operating of the equipment. It is also necessary to regulate the wash water added to U04-

CG01 and U04-CG02 to achieve effective washing of the solids in the slurry. The final

product, stored in U04-ST04 with a capacity of 100 m3, has a solid content of 35% and this

product is sent to NUFCOR which process the ADU to produce U3O8.

3.10. Innovations

The design procedure proposed by Douglas (1988) creates a platform for the development

of creative solutions for specific design problems. All the proposed solutions are

economically evaluated to ensure that profitability is not compromised and the impact on the

environment and working conditions are also studied. Table 3.16 shows the innovative

steps taken in the process development.



Table 3.16: Innovations

Innovations Advantages

Use of nitric acid instead of MnO2 for

leaching oxidant.

- Nitric acid provides faster kinetics

- Less nitric acid is used due to the

auto-catalytic effect which reduces

the raw material costs.

Use of existing leaching pachucas. Due to

the faster kinetics, the increased feed can

still be processed in the existing columns.

- No additional pachucas needs to be

constructed which means less capital

is needed for the expansion project.

Use of existing leaching pachuca for

neutralization because not all the pachucas

are used for the new leaching process.

- No additional construction is needed

for the neutralization plant.

- All the existing puchacas are used

therefore decommissioning is not

necessary.

Construction of additional CCD train instead

of decommissioning of existing train. The

existing CCD equipment will not be able to

handle the additional feed, therefore an

additional CCD train is built and the feed

stream is divided between the two trains.

- The existing CCD equipment is still

used therefore no decommissioning

is necessary.

- This setup will be able to process

more feed than required for this

project which allows further increase

in ore processing if necessary.

Use of existing adsorption ion exchange

columns for fixed-bed ion exchange.

- No additional ion exchange columns

needs to be built, therefore less

capital is needed.

Use of after-settlers in the solvent extraction

process.

- The after-settlers provide additional

residence time to allow phase

separation and this reduces solvent

loss. The solvent used is extremely

expensive thus it is critical to

minimize solvent loss.



Use of four stripping mixer-settlers in the

solvent extraction process, instead of three.

- The pH control is simplified by adding

an additional mixer-settler. The pH

control is difficult because the recycle

stream from the precipitation is too

high. Therefore the first settler is

used to allow the solution pH to reach

a steady state which simplifies the pH

control in the remaining settlers.

Use of bio-organisms to remove the nitrate

ions in the waste water.

- Removes the excess nitrate ions in

the waste water which is dangerous

to the environment.

Construction is done over a three-year

period of which the first year the existing

plant will still operate at normal capacity

while the additional equipment is build. The

second year the plant will shut-down for

construction and the final year the plant will

be operated at 50% of the new capacity.

- Several of the new equipment can be

constructed while the old plant is still

operating, therefore production is not

lost.

- Because various existing equipment

is used and only a few facilities needs

to be decommissioned, the

construction period is short, therefore

production loss is minimized.


108 Chapter 4: Detail design

Chapter 4: Detail design

The purpose of the solvent extraction (liquid-liquid separation) process is to purify and

concentrate the uranyl sulphate ions. The principle of the solvent extraction process is

based upon immiscible properties of aqueous and organic liquids. The existing solvent

extraction section on the South Uranium Plant (SUP) is out-dated and should be upgraded.

The detail design for the solvent extraction process includes the following considerations:

• Choice or type of system

• The kinetics and thermodynamics of the system.

• A detail chemical design.

• The start-up and shut-down procedures.

• Optimization of the process which includes sensitivity analysis.

• Mechanical aspects that should be considered for the process.

The solvent extraction process consists of four separate stages which include: extraction,

scrubbing, stripping, and regeneration. In the extraction stage the uranyl sulphate ions are

removed from the aqueous phase into the organic phase. In the scrubbing stage the organic

phase is cleansed from impurities and in the stripping stage the uranyl sulphate ions are

transferred from the organic phase to aqueous phase. During the stripping stage the

concentration of the uranyl sulphate ions is increased. The regeneration stage removes any

impurities entrained in the organic solvent from where the solvent is recycled to the

extraction stage.

4.1. Choice of system type

The different process alternatives for the solvent extraction section are discussed in Section

2.5.3.1. The first choice is the type of extractant that should be used for the system,

followed by the composition of the solvent. Alamine® 336 is used as extractant for this

system which has been used for the past 20 years, and is a tried and tested extractant for

the SUP process. The current solvent composition (volume %) is also used which contains:

93% kerosene, 5% Alamine® 336, and 2% isodecanol.



There is a choice which should be made for the equipment between mixer-settlers and

pulsed columns. Mixer-settler equipment, shown in Figure 4.1, was chosen above pulsed

columns to ensure easier pH control for the stripping stages of the process (Law & Tod: 4).

The current operators are experienced in the operation of mixer-settler equipment therefore

operating optimization will need less time and fewer resources will have to be spent on

training. If pulsed columns is used the operating personnel will have to be engineers,

making the labour cost extremely high. In the case of unforeseen plant disturbances all the

solvent in the pulsed columns will be lost which is economically unfavourable as the solvent

is very expensive (Watson, 2009).

Figure 4.1: Representation of mixer-settler equipment

The operating choices for the flow of the mixer-settler equipment are cross-current and

counter-current flow configuration of the organic and aqueous phases. Counter-current flow

is used for this system due to higher overall efficiency. Counter-current flow allows for the

best compromise between high recovery and good separation for the process.



4.2. Kinetics and thermodynamics

The mixer-settler equipment, as seen in Figure 4.1, consists of two separate compartments,

i.e. mixing and settling. In both these compartments the kinetic and thermodynamic

parameters are vital to the design of the mixer-settler equipment. The kinetics for the

extraction of uranium from the aqueous phase to the solvent phase containing Alamine® 336

is extremely fast, therefore equilibrium parameters are more important for the mixing

compartments (Mackenzie, 1997: 11).

To describe the equilibrium parameters for the extraction of uranium, loading isotherms are

needed. Loading isotherms are described by the equilibrium distribution of uranium in the

aqueous and organic solvent phases. It is reported by Stönner & Wiesner (1982) that the

loading isotherm for the extraction stage of solvent extraction is modelled by Equation 4-1.

[ ][ ]

[ ] [ ]−=

+ + +

aq1org

2 3 2 4 4aq aqaq

UU k

U k Cl k H SO k (4-1)

All the concentrations are in g/L and the concentration of uranium is in U3O8. The constants

k1 to k4 are fitted parameters with k1 being the maximum uranium loading capacity of the

organic phase. Modelling using Equation 4-1 agrees with data supplied by the Metallurgy

Division of the South African Atomic Energy Board (Stönner & Wiesner, 1982: 97).

The loading isotherm for the stripping stages can be found in Morais & Gomiero (2005) and

is for chloride free uranium ores. This loading isotherm is displayed in Figure 4.2.



Figure 4.2: Stripping loading isotherm with (NH4)2SO4

Using the above mentioned equilibrium data, it is possible to design the required number of

stages for effective separation. Due to the fast kinetics for the uranium transfer between the

phases a short residence time in the mixing compartments is required.

The main objective of the settlers is to achieve efficient separation between the aqueous and

organic phases. Stönner & Wiesner (1982) reported that the settling kinetics can be

modelled by two consecutive steps. The first step is a volume-controlled reaction which

describes the coalescence of small droplets to large droplets. In the second step which is

area-controlled, the large droplets enter its original phase. The kinetics for these steps is

given in Equations 4-2 and 4-3 (Stönner & Wiesner, 1982: 98).

= −1 1dV Vk.F.Hdt V (4-2)

= −2 2dV Vc.Fdt V (4-3)



Equation 4-2 represents the kinetics of the first step where V1 is the volume of dispersion

contains small droplets, F is the settling area, H is thickness of the dispersion band and V is

the total volume of dispersion. Equation 4-3 represents the kinetics of the second step

where V2 is the volume of dispersion containing large droplets and c is the velocity of these

droplets (Stönner & Wiesner, 1982: 98).

4.3. Detail chemical design

The following procedure is followed to do a detail chemical design on the solvent extraction

unit. This unit consists of four sections as mentioned before. Each of these stages consists

of a number of mixer-settler units connected using counter-current flow. The first step in the

design procedure is to determine the amount of mixer-settler units in each section. The next

step is to size the equipment, starting with the mixers. After the mixers are sized, the

settlers, pumps and tanks are sized. The assumptions made for the chemical design are the

following:

• Uranium recovery of the extraction section is 98%.

• Extraction and stripping stage efficiency of 98%.

• Mixer agitators work in turbulent flow regime.

• Height of the dispersion layer is directly proportional to the volumetric flow of the

droplets.

• Initial volumetric fraction of small droplets in the dispersion is 100%.

4.3.1 Number of stages

Using the extraction and stripping loading isotherms together with a basic mass balance, the

McCabe-Thiele method is applied to determine the number of stages. Accurate isotherms

are required for this step, which is stated in Section 4.2 (see Appendix C).

The isotherm for the extraction section is obtained from Equation 4-1 and the constants were

manipulated to obtain a similar trend. The given constants, reported by Stönner & Wiesner

(1982), and manipulated constants for Equation 4-1 is shown in Table 4.1.



Table 4.1: Extraction loading isotherm constants

Constant Value reported Manipulated value

k1 6.772 6.772

k2 0.77218 0.772

k3 0.02864 0.002864

k4 2.51 0.5

The validation of the manipulated variables is shown in Appendix C. The design method

using McCabe-Thiele is discussed in detail in Appendix C. From Figure C.4 and C.5 the

exiting uranium loading is obtained and shown in Table 4.2.

Table 4.2: Results from McCabe-Thiele method

Stage Organic loading g U3O8/L Aqueous loading g U3O8/L

Extraction

1 3.5918 0.95

2 1.05 0.15

3 0.3 0.07

Stripping

1 11.95 1.4

2 4.15 0.7

3 1.7 0.28

The organic phase fed to the extraction section has uranium loading of 0.25 g U3O8/L and

exits at 3.5918 g U3O8/L, while the oraganic phase fed to the stripping section has uranium

loading of 3.5918 g U3O8/L and exits at 0.28 g U3O8/L. Next the mixers are designed to

obtain the desired efficiency.

4.3.2 Mixer

The design considerations for the mixer are residence time, impeller rotary speed, flow

regime, impeller tip speed, produced head and mixer box dimensions. The residence time

for all mixers was assumed from literature to be 2 minutes (Mackenzie, 1997: 11). The

results obtain using the design philosophy in Appendix C is shown in Table 4.3.



Table 4.3: Mixer box design results

Specifications Value

Hydraulic efficiency 0.23

Impeller diameter 1.1 m

Impeller rotary speed 70 rpm

Impeller tip speed 4 m/s

Mixer box diameter 1.4 m

Mixer box height 1.7 m

Mixer box liquid volume 2 m3

Nh 2.46

Np 0.5493

Nq 0.01048

Produced head 2 m

Residence time 2 min

Reynold number 1.7 x 106

Total volumetric flow (10% over design) 59.6

From Table 4.3 it is seen that turbulent flow is achieved which allows for adequate mixing

between the different phases. The produced head is greater than the height of the mixer

box, which will ensure continuous flow into the settler vessel. The residence time is 2

minutes which ensures sufficient mixing, and will therefore allow for high stage efficiency.

4.3.3 Settler

The settling kinetics used in this design is derived from Equation 4.2 and 4.3 and shown in

Appendix C. The kinetics were solved using Polymath® and sensitivity analyses were done

on the dispersion layer velocity, width of the setter vessels and construction cost. The

design constraints are identified as the following:

• Height of the vessel should not exceed 1.5 m, to reduce the mixer head required.

• The width of the vessel should not exceed 2 m, to ensure adequate dispersion

distribution into the vessel.

• Final dispersion layer height should not exceed 1 mm, to reduce the solvent loss.

• A realistic residence time should be achieved, to ease operability.



Using the settling kinetics and above mentioned constraints, the settler vessel dimensions

are sized for the extraction and stripping sections. The designed dimensions are shown in

Table 4.4.

Table 4.4: Designed dimensions for the settler vessels

Dimension Value

Extraction

Height 1.5 m

Width 2 m

Length 5.5 m

Residence time 12.5 min

Initial dispersion height 1.1 m

The sizes of the vessels for the extraction and stripping sections differ due to the difference

in aqueous to organic feed ratio. The results in a lower flow for the stripping section,

therefore, smaller settler vessels are required for the stripping section. However, since the

material of construction is vacuum infused fibreglass, a mould is used to construct the

settling vessels. The construction cost for this manufacturing method consist of the price for

each mould and construction materials used, therefore, using one standard vessel size will

optimize construction cost. Due to the smaller vessels required for the stripping section, it is

decided that the size for both the stripping and scrubbing settler vessels are exactly the

same as for the extraction settling vessels.

4.3.3 Pipe sizing

The pipes are sized to achieve a velocity flow less than 1 m/s to reduce the probability of a

fire hazard caused by statistic discharge. The design methodology used to sizes the pipes

are shown in Appendix C and the results given in Table 4.5.



Table 4.5: Solvent extraction pipe sizes

Over-designed flow Pipe inner diameter Flow velocity

(m3/hr) (m) (m/s)

Eluate flow 28.4 0.115 0.76

Organic flow 31.2 0.115 0.84

Stripping flow 8.7 0.065 0.73

Scrubbing flow 3.1 0.04 0.69

All the chosen pipe inner diameters are standard pipe sizes (converted from inches) to avoid

the cost for custom made pipes.

4.3.4. Pump sizing

The design of the solvent extraction section requires 11 pumps. The pumps are required to

transport the liquid from the respective storage tanks to the mixers. The design

methodology used for the pumps are discussed in Appendix C and the results obtained for

the pump sizing is shown in Table 4.6.

Table 4.6: Pump sizing results

Pump equipment name Head required (m) Liquid flow (L/s)

U03-PM01 4 7.2

U03-PM02 17 6.8

U03-PM03 4 7.2

U03-PM04 3.5 7.9

U03-PM05 4.5 1.6

U03-PM06 3.5 1.6

U03-PM07 4.5 2.2

U03-PM08 2.5 2.2

U03-PM09 5 1.6

U03-PM10 4.5 1.6

U03-PM11 6.5 1.6

Once a pump supplier is contracted, the pump curves are used to determine which pumps to

use. When choosing a pump, using the pump curves, it is important to use a pump with high

efficiency at the desired head and liquid flow. It is also important to calculate the NPHS (net

positive head suction) required for the pump and compare it with the achieved PHS (positive

head suction) for the system.



In case of a pump malfunction or maintenance on the pump, it is important to design a back-

up pump system. The system should always be on standby to ensure continuous operation

in case of pump failure. The configuration of the back-up pump (used for all the pumps) is

shown in Figure 4.3.

Figure 4.3: Back-up pump configuration

A control valve is placed before and after each pump configuration to open and close liquid

supply as needed. Routine maintenance needs to be done weekly, which includes oil

change and re-greasing of the pump motors. Weekly inspections are done to check the

gaskets for any leaks, while vibration and temperature tests must be done semi-annually.

Due to maintenance the pumps are alternated each week to ensure continuous operation of

the plant.

4.3.5. Control valve sizing

For the preliminary feasibility study it is important to determine the pressure drop (∆P) across

the valve, volumetric flow (Q) and the valve coefficient (CV). It is also important to determine

the inherent valve characteristics required to ease upgrade and controllability of the plant.

The inherent valve characteristics chosen for all the valves are equal percentage valve

characteristics. The equal percentage characteristics is chosen due to the forgive nature of

the valve. If the pumps are incorrectly design an equal percentage valve will show better

valve characteristics which will not be seen for linear or quick opening valves. Due to this

fact, the equal percentage valve is used when upgrading a system (Svrek et al, 2006:34).



The volumetric flow is calculated in the mass balance and is therefore a known variable.

The pressure drop is calculated using Equation C-22 and that leaves the valve coefficient

which is calculated using Equation C-23. The resulting pressure drop and valve coefficients

are shown in Table 4.7.

Table 4.7: The calculated pressure drop and valve coefficients

∆PCV (kPa) CV (m3/hr.bar0.5) CV (gpm/psi0.5)

U03-FCE01 17.522 7.16687 8.37251

U03-FCE02 19.4698 5.85162 6.836

U03-FCE03 19.4698 5.85162 6.836

U03-FCE04 17.522 7.16687 8.37251

U03-FCE05 19.222 5.85504 6.84

U03-FCE06 19.6822 1.2803 1.49568

U03-FCE07 19.672 2.25463 2.63392

U03-FCE08 18.772 2.30805 2.69632

U03-FCE09 18.772 1.15402 1.34816

U03-FCE10 19.6757 1.77874 2.07797

U03-FCE11 17.522 7.16687 8.37251

U03-FCE12 19.672 0.45093 0.52678

U03-FCE13 19.672 0.45093 0.52678

U03-FCE14 19.672 0.45093 0.52678

U03-FCE15 18.7257 1.8233 2.13002

U03-FCE16 17.522 7.16687 8.37251

U03-FCE17 17.522 1.19448 1.39542

U03-FCE18 17.522 2.38896 2.79084

U03-FCE19 19.672 6.40316 7.48032

U03-FCE20 17.522 6.1325 7.16414

To chose a valve once a supplier is contracted, the calculated valve coefficient shown in

Table 4.7 is used. It is also very important to size the valve according to the pipe diameter to

ensure a small additional pressure drop which will lead to superior valve characteristics.



4.4. Start-up and shut-down procedures

Once a solvent extraction section is designed and built, it is important to have a proper

start-up procedure. During the start-up procedure for the commissioning period, all process

equipment and instrumentation are tested to ensure the desired performance. It is also

important to have a shut-down procedure in case of an emergency or upstream or

downstream process units. The start-up procedure after such a shut-down is less intense

and time consuming, since most of the equipment testing was done during the

commissioning phase.

4.4.1. Commissioning of process unit

The commissioning of the solvent extraction section can be categorised into several steps

during which specific procedures are followed to achieve successful operation of the process

unit. The different commissioning steps are given chronologically below.

Dry inspection

When the solvent extraction section is constructed, all process instrumentation is manually

inspected for defects and other problems. A list of the more important inspections is given

below.

• Valves

• Connections at all flanges.

• Clear all obstructive objects from mixers and settlers.

• Ensure all pH probes are installed.

• Ensure all conductivity meters are installed.

• Ensure all flow meters are installed.

• Electrical power to all equipment and instrumentation.

• Fire extinguisher equipment.

The dry inspection is a tedious step in commissioning of a process unit but is of utmost

importance to prevent equipment damage at start-up. When these inspections are done the

process unit is ready to receive liquid feed.



Wet inspection

In this step the solvent extraction section is flooded with demineralised water. It is important

to make sure that all the valves are in the open position before the wet inspection is started,

to avoid high differential damage to the equipment due to water hammering. It is necessary

to make sure that the control system is not active during this step. When all the equipment is

at maximum capacity several inspections are carried out.

• Flow direction of the pumps.

• Rotational direction of the mixer impellers.

• Pipe and vessel leaks.

• Operation of the control equipment.

• Accuracy of the pH probes.

• Accuracy of flow meters.

• Operation of the conductivity meter.

• Effective working of weirs in the mixers and settlers.

It is also important to test the control system for the process unit. Once these inspections

are complete, the solvent extraction section is ready to receive the organic feed.

Preparation for operation

When the solvent extraction section is flooded with water and all required inspection are

completed to satisfaction, the organic solvent is allowed onto the section. The organic

solvent is introduced into the system through its designed flow. Since the density of the

organic solvent is lower than that of the water still being fed to the process unit, the top

sections of the settler vessels will be filled with organic solvent. Approximately 66 m3 of

solvent is required before the settler vessels are operated at the designed levels.

From here the control system is activated to achieve the correct flow rates and to start the

pH control on the stripping section. When everything is in place, the water in the solvent

extraction section is replaced with the specific streams, first the stripping agent for the

stripping section and finally the eluate in the extraction section to introduce uranium into the

system. As soon as the first uranium is produced in the OK-liquor the solvent extraction

process unit is signed over to the operating personnel who will further optimise it.



4.4.2. Shutdown procedures

There are two possible shut down procedures for the solvent extraction section, one for

when the settlers are required to be emptied and one where only the flows in and out of this

process unit is stopped. The latter procedure is used more often since it holds a lower

safety hazard while effectively halting operation by requiring no feed from the previous

section and producing none for the next. Since counter-current flow is used through the

solvent extraction section, if one mixer-settler needs to be shut down, then all mixer-settlers

are shut down.

During the shut down procedure mentioned first, the pumps to the section is stopped while

the drain valves on the settler vessels are opened to drain the contents of these vessels.

This procedure is dangerous since the highly flammable organic solvent runs freely through

a drain. The organic solvent then ends up in a sump where it is preserved and the aqueous

phase is separated from it. If this shut down procedure is used it is important to follow the

start up procedure of a wet inspection followed by the preparation for operation procedure.

The simpler shutdown only halts the feed from the ion exchange and to the ADU

precipitation. This is achieved by stopping all the pumps while keeping the impellers in the

mixers rotating. This causes the internal recycle streams to be continuously recycled while

no organic or aqueous phases transfers between the different mixer-settler stages. The

pump around through the internal recycle may be carried out for long periods supplying

enough time to repair problems and complete inspections on the plant. If the shut-down is

expected to last longer than 2 days, the pumps and mixers are turned off. The start-up

procedure after the above mentioned shut-down procedures are to firstly start the mixers to

allow for internal recycle flow. When adequate internal flow is achieved, the pumps are

started, which will result in normal operation.

4.4.3. Emergency shutdown procedures

Emergency shutdown procedures are set in place to ensure plant and personal safety in

case of unforeseen circumstances. Three possible disturbances are evaluated and a

shutdown procedure is suggested for each.



Low flow or high flow

Disturbances upstream can lead to unwanted high or low flow velocities. These unwanted

flows can cause phase inversion which results in solvent lost which is economically

unfavourable as solvent is very expensive. If high or low flow occurs, the following

procedure should be followed:

• The pumps should be turned off, while the mixers are still running. Switching of the

all pumps will stop any feed to enter the system which will reduce the probability of

phase inversion. Internal recycle, due to the mixers, will ensure that the liquids inside

the settler vessel are continuously moving.

• The situation should be evaluated to determine the correct action that should be

done to restore the operability of the process.

• If the proposed corrective action is time consuming, the mixers can be turned off to

save electricity and therefore decreasing operating costs.

• The corrective action should then be implemented and the situation should be

monitored closely to ensure that the problem is solved and normal operating

conditions are resorted.

No flow or low level

No flow in the system or low levels in the equipment will cause a solvent extraction section

trip, which can lead to a plant-wide shut down. These low levels and no flow rate will also

cause phase inversion which will lead to solvent loss. In case of a plant trip due to no flow or

low level the following emergency shutdown procedure is followed:

• All the pumps and mixers should be shut down to ensure equipment safety and avoid

loss of solvent due to phase inversion.

• The situation should be evaluated to determine the corrective action.

• The corrective action should then be implemented and the situation should be

monitored closely to ensure that the problem is solved and normal operating

conditions are resorted.



Maintenance

On any plant, maintenance is always required to repair equipment and instrumentation.

Maintenance usually involves grinding or welding which is a major fire risk on solvent

extraction section. It is therefore crucial to have a detailed shut down procedure for when

maintenance is required on the solvent extraction section. The following emergency

shutdown procedure should be followed:

• The entire solvent extraction system should be drained to the sump except for the

storage tanks. The tanks are isolated from the system using the different control

valves used for flow control.

• All the streams and equipment should be flushed with potable water to wash away

the flammable organic solvent.

• If possible, the damaged equipment pieces should be removed from the demarcated

area and repaired outside the dead zone (discussed in Chapter 6).

• Finally, when the repair work is completed, the wet inspection start-up procedure

should be followed and the preparation for operating should be conducted.

4.5. Mechanical aspects

The mechanical design of the mixer-settler equipment is a direct result of the chemical

design with some extra mechanical considerations. In this section the mechanical drawings

created in SolidWorks® 2009 SP2.1 are given and discussed. As mentioned before, the

material of construction for the mixer and settler vessels is vacuum infused fibre glass, while

the extra components such as the impeller, inlet manifold and picket fence are constructed

from stainless steel 316L. The mechanical drawings consist of three pages:

• Page 1 of 3: Overview and assembly of equipment.

• Page 2 of 3: Mixer equipment with important dimensions.

• Page 3 of 3: Settler equipment with important dimensions.



In the assembly drawing of the mixer-settler equipment (Page 1 of 3) the mixer and settler

vessel is shown together with the basic equipment. These are shown in an assembled and

exploded view. A list of this equipment follows.

• Motor

• Mixer vessel

• Mixer lid

• Impeller

• Inlet manifold

• Settler vessel

• Picket fence

The assembly of these parts are indicated on the mechanical drawing. The mixer lid is also

constructed from vacuum infused fibre glass. More detailed drawings of the mixer and

settler equipment are given in pages 2 and 3 of the mechanical drawings.

The mixer equipment shown on Page 2 of 3 includes the mixer vessel, mixer lid, impeller,

inlet manifold and the representation of the electric motor. The mixer vessel has an inner

diameter of 1.4 m with a height of 1.7 m. The bottom part of the vessel slopes down at an

angle of 30° to meet the inlet manifold. The liquid outlet of the vessel starts 0.3 m from the

top with a height of 0.2 m and width equal to the inner diameter of the mixer vessel (1.4 m).

A representation of the impeller is given in Figure 4.4.

Figure 4.4: Pumper mixer Impeller



The impeller is designed with a diameter of 1.2 m while the mixer vessel is equipped with

baffles extending 0.05 m as indicated on the mechanical drawing. This leaves a space of

0.05 m between the impeller and baffles. The impeller used for this design is a product of

MC Process, and the detail mechanical specifications are confidential. The mixer vessel is

fitted to the inlet manifold with a diameter of 0.152 m at the bottom. The mixer and settler

vessels are attached with a flange plate at the liquid outlet.

In order to extent the life-time of the pH probes used in the settlers, it is important to ensure

that it does not come in contact with the organic solvent. To reduce lag time in the control of

pH the measurement should be done in the mixer. To overcome the problem a small

external pH box is attached to the mixer vessel which causes separation. This allows for pH

measurement without damage to the expensive pH probes. Figure 4.5 is a schematic

representation of the pH box.

Figure 4.5: pH box for measurement

The settler equipment is shown on Page 3 of 3 of the mechanical drawings. The settler

vessel has a total length of 8.3 m which consists of the connection to the mixer vessel, the

settling compartment and the aqueous outlet compartment. The connection to the mixer

vessel has a directional length of 0.6 m. The settler compartment is defined as the length

between the liquid inlet and the aqueous weir and has a length of 5.5 m, as designed. The

aqueous outlet compartment is separated from the settling compartment by the bottom weir

and is 1.8 m in length as indicated in the mechanical drawing. It is also important to note

that the settler has a breathable lid to reduce solvent loss due to evaporation.



The mechanical drawing on Page 3 of 3 also shows the organic phase weir, aqueous phase

weir and the overflow. Each of these weirs is equipped with pipes to enable the different

streams to exit the settler vessel. It should be noted that the aqueous outlet weir is 0.14 m

lower than the organic outlet weir to account for the difference in liquid levels due to the

density difference between the organic and aqueous phase. Picket fences are used in the

settling compartment to enhance phase separation. As indicated on the mechanical

drawings, these picket fences fit into the picket fence slots on the settler vessel. Three

picket fence slots are situated 1 m apart in the settler vessel which allows for adjustment of

the picket fence position to obtain optimum performance.

These mechanical aspects are a preliminary guideline for the mechanical design of the

mixer-settler equipment. These mixer-settlers are used in the extraction, scrubbing,

stripping and regeneration sections of solvent extraction. The given specifications result

from the chemical design and ensure adequate sizes for an over-designed solvent extraction

process.



Mechanical drawing: Page 1


130 Chapter 5: Techno-economical evaluation

Chapter 5: Techno-economical evaluation

The purpose of any chemical plant is to make a profit and to increase the value of the initial

investment into the venture. The techno-economical evaluation is important to ensure that a

profit is obtained after the proposed lifetime of the project; this is done by certain economical

analyses. The considerations for the techno-economical evaluation are:

• Equipment cost.

• Fixed and working capital.

• Fixed and variable cost.

• Revenue from product and by-product sales.

Due to the fact that this project is an upgrade for an existing plant, some of the existing

equipment and facilities will remain in the process which will decrease the equipment and

overhead plant cost involved in this evaluation. This techno-economical evaluation consists

of:

• Assumptions and definitions.

• Estimation of capital, operating cost and income.

• A cash flow analysis.

• Economic sensitivity analysis.

• Recommendations for profitability.

5.1. Definitions and assumptions

The economic evaluations done in the following sections are better understood if the

definitions and assumptions are known. The different definitions and assumptions are

described in this section.



5.1.1. Definitions Fixed capital investment (FCI) – The investment needed to purchase or manufacture and

install all the process equipment with all components and necessary plant facilities (Peters et

al, 2004: 233).

Working capital – The capital investment necessary for the operation of the plant which

include the money needed for raw materials, taxes, monthly expenses such as salaries and

wages and accounts that need to be paid (Peters et al, 2004: 233).

Fixed costs – The fixed costs include all the expenses which are independent of the

production rate of the product as well as the plant overhead costs and general expenses

associated with the management of the company (Peters et al, 2004: 262).

Variable costs – The variable costs is all the expenses associated with the manufacturing of

the product and is depended on the operability of the plant (Peters et al, 2004: 262).

Revenue from sales – The sale of the products produced by the plant is known as the

revenue and is calculated as the sum of the unit price of each product multiplied by the rate

of sales (Peters et al, 2004: 258).

Income taxes – Income taxes are paid on a corporate-wide basis with the total gross profit

being the taxable income of a corporation. The tax rate is subject to change as it is

depended on the taxable income (Peters et al, 2004: 304).

Discounting – The calculation used to determine the present worth of a future amount using

a discount rate which can be calculated with Equation 5-1. In Equation 5-1, i is the interest

rate and N is the number of years (Peters et al, 2004: 298).

( )−= +NDiscount factor 1 i

(5-1)

A country’s central bank also has a discount rate at which it charges commercial banks for

loans to meet the temporary shortage of funds.



Inflation – The percentage change in consumer prices of goods and services compared with

the previous year’s prices, and this is regardless of the time value of money (Peters et al,

2004: 290).

The financial resources required for a project is limited and should be used in an appropriate

and efficient manner. Different alternatives are available to determine the most efficient use

of these resources. To evaluate these alternatives, methods are used to calculate the

profitability of the project and these methods will be discussed.

Return on investment (ROI) – This method does not consider the time value of

money. The method calculates the annual return on an investment as percentage

per year, by dividing the annual net profit by the total capital investment. It is

recommended that the average ROI over the entire project life is calculated using

Equation 5-2, to obtain a better representation of the profitability (Peters et al, 2004:

323).

( ) ( )=

=−

=∑

∑

N

p,jj 1

N

jj b

1 NNROI

F (5-2)

In Equation 5-2 N is the evaluation period, Np,j is the net profit in year j, -b the year in

which the first investment is made in the project with respect to zero as the startup

time, and Fj is the total capital investment in year j.

Payback period (PBP) – This method does not consider the time value of money

and is a method used to calculate the time needed for the cash flow to equal the

original fixed-capital investment. This payback period is calculated using Equation 5-

3 in which V is the manufacturing fixed-capital investment, Ax the nonmanufacturing

fixed-capital investment and (Aj)ave the average cash flow (Peters et al, 2004: 324).

( ) ( ) ( )=

+ += =

∑x x

Nj avej

j 1

V A V APBPA1 AN

(5-3)

A project’s PBP should be less than or equal to the reference value calculated using

Equation 5-4 (Peters et al, 2004: 324).



=+ar

0.85PBP 0.85m N (5-4)

Equation 5-4 is the PBP obtained from the minimum acceptable rate of return, mar

expressed as a percentage per year (Peters et al, 2004: 321).

Net present value (NPV) – This method consider the time value of money. Using the

present worth of all cash flow and all capital investments in Equation 5-5, the net

present worth can be calculated (Peters et al, 2004: 327).

( )( )= =−

= − − −Φ + + − ∑ ∑N N

cf,j j oj j j j v,j jj 1 j b

NPV PWF s c d 1 rec d PWF F (5-5)

In Equation 5-5, PWFcf,j is the selected present worth factor for the cash flow, sj is the

value of sales, coj is the total product cost not including depreciation, recj is the

dollars recovered from the working capital and sale of physical assets, PWFv,j the

appropriate present worth factor for investments occurring, and Fj the total

investment. All these variables are values for year j.

Internal rate of return (IRR) – This method is also known as the discounted cash

flow rate of return (DCFR) and consider the time value of money. This method

calculates the return using all investments and cash flows which are all discounted.

Equation 5-5, which calculates the net present value, is set equal to zero and the

discount rate is solved (Peters et al, 2004: 328).

Depreciation – Physical facilities decreases in value with time in terms of physical

depreciation and functional depreciation. The causes of physical depreciation include wear

and tear, corrosion, and accidents while functional depreciation is caused by technological

advances, making an existing property obsolete (Peters et al, 2004: 307).



5.1.2. Assumptions

It is necessary to use certain assumptions to simplify the calculations. It is a common use to

assume fixed amounts for the different rates; although rates are subject to change, these

changes cannot be accounted for in these analyses. The following assumptions were used

in this techno-economic evaluation.

• Tax rate of 28% (The Low Tax Network, 2009);

• Discount rate of 11% (Photius, 2009);

• Inflation rate of 6.4% (Viljoen, 2009);

• Rand/dollar exchange rate of 7.42 R/dollar (X-rates, 2009);

• The construction of the upgraded equipment for this plant will be done over a period

of two years. For this economic evaluation the assumption was made that the plant

will be in operation for 20 years therefore the evaluation was done for a 22-year

period.

• During the first year of construction the plant will be operated at the usual capacity of

producing 624 ton of ADU per year while the additional CCD equipment is being

constructed. The second year of construction the plant will be shut down to conduct

necessary piping changes for the new process and construction of the new solvent

extraction section.

• The first year of operation after the construction period the plant will run at 50% of the

new capacity to allow for operational troubleshooting.

5.2. Estimation of capital, operating cost and revenue

Whenever a new project is launched, sufficient capital is required to construct the equipment

and facilities needed to produce a product which will sell at a profit. This capital investment

should be calculated accurately to ensure the maximum rate of return on the initial

investment. It is also important to estimate certain operating costs that will be required to

operate the plant throughout the project lifetime. The revenue produced from sales should

surpass the initial capital investment and operating costs to ensure economic feasibility of

the project.



5.2.1. Capital investment

Different methods can be used to calculate the total capital investment that is needed for a

project. The percentage of delivered-equipment cost method is based on the determination

of delivered-equipment cost and other direct and indirect costs. This method uses

Equation 5-6 to calculate the total capital investment (Cn) (Peters et a.l, 2004:250).

( )= + + + +∑n 1 2 nC E 1 f f ... f (5-6)

In Equation 5-6 f1 through fn are the multiplying factors for the specific direct and indirect

costs and E is the total purchased cost of the equipment. The values of these multiplying

factors are given in Table 5.1 on the following page. The working capital is 15% of the total

capital investment. Because this is an expansion project for the South Uranium Plant,

buildings and services already exist which will lower the total capital investment needed by a

factor of 0.8.

For the extraction of uranium the solid-fluid processing plant factors are used to calculate the

total capital investment with the delivered-equipment cost method. These calculations and

assumptions are given in Appendix D, with the following results:

• Fixed capital investment (FIC) = R 720 million.

• Working capital = R 126 million.

• Total capital investment = R 847 million.

After calculating the total capital investment, it is important to estimate the operating cost

required to run the plant.



Table 5.1: Multiplying factors used in the delivered-equipment cost method

Fraction of delivered equipment cost for

Solid

processing

plant

Solid-fluid

processing

plant

Fluid

processing

plant

Direct cost

Purchased equipment (E) 1 1 1

Delivery 0.10 0.10 0.10

Purchased equipment installation 0.45 0.39 0.47

Instrumentation and controls (installed) 0.18 0.26 0.36

Piping (installed) 0.16 0.31 0.68

Buildings (including services) 0.25 0.29 0.18

Electrical system (installed) 0.10 0.10 0.11

Yard improvements 0.15 0.12 0.10

Service facilities (installed) 0.40 0.55 0.70

Total direct plant cost 2.69 3.02 3.60

Indirect plant cost

Engineering and supervision 0.33 0.32 0.33

Construction expenses 0.39 0.34 0.41

Legal expenses 0.04 0.04 0.04

Contractor’s fee 0.17 0.19 0.21

Contingency 0.35 0.37 0.44

Total indirect plant cost 1.28 1.26 1.44

Fixed-capital investment 3.97 4.28 5.04

Working capital 0.70 0.75 0.89

Total capital investment 4.67 5.03 5.93



5.2.2. Operating cost

The operating costs considered for this project is for the entire plant, and not just for the

upgraded sections. The operating costs consist of variable production cost and fixed

charges which include plant overhead costs and general expenses. The factors used to

calculate the operating costs and the amounts calculated is showed in Table 5.2 with

detailed calculations given in Appendix D.

Table 5.2: Operating costs

Fraction Fraction basis Calculated amount (R million per

annum)

Raw materials 116.39

Operating labor 4.96

Operating supervision 0.15 Operating labor 0.74

Utilities 25.87

Maintenance and repairs 0.07 Fixed-capital investment 50.45

Operating supplies 0.15 Maintenance and repairs 7.57

Laboratory charges 0.15 Operating labor 0.74

Royalties 0.04 Revenue from sales -

Catalyst and solvents 4.86

Total variable cost 211.59

Taxes (property) 0.02 Fixed-capital investment 14.42

Financing (interest) 0.105 Fixed-capital investment 75.68

Insurance 0.01 Fixed-capital investment 7.21

Rent 0 Fixed-capital investment -

Plant overhead costs 0.5 Sum of operating labor,

supervision and

maintenance

28.08

General expenses 0.2 Operating labor 0.99

Total fixed cost 126.37

Total product cost 337.96

From Table 5.2 it is seen that the fixed cost is R 126 million per annum while the total

variable cost is manipulated to obtain the variable cost per ton product. At a production rate

of 1068 ton ADU per year, the variable cost amounts to R 198 000 per ton ADU.



5.2.3. Revenue

Up to now the necessary capital investment has been calculated, which is only a part of the

complete cost estimation. The revenue generated from the plant operation also plays an

important role in the complete cost estimation. The product of the South Uranium Plant is

ADU with the by-product of 0.5 g gold per ton ore. The commercial and proposed sale price

of the product and by-product are given in Table 5.3 with the annual revenue produced.

These calculations are also supplied in Appendix D.

Table 5.3: Revenue from product and by-product sales

Product Commercial selling price (R/ton)

Proposed selling price (R/ton)

Annual revenue (R million per annum)

ADU 743 786.41 198 079.21 211.55

Gold 254 552 770.00 127 276 385.00 135 931.18

The assumptions made for the proposed selling price of the ADU and gold are:

• AngloGold Ashanti only receives half of the commercial selling price of ADU.

• South Uranium Plant receives 50% of the gold extracted.

The selling price per unit is calculated from data received on 26 October 2009 from Mr.

William Manana at the South Uranium Plant with a spot price of 45.50 $/lb for uranium and

254 552.77 R/kg for gold.

5.3. Cash flow analysis

The total capital investment, operating cost and revenue from sales calculated are used to

perform a cash flow analysis for the South Uranium Plant. The proposed lifetime of the plant

is 20 years with a construction period of two years to construct the new counter-current

decantation and solvent extraction sections and to upgrade the required piping. After the

two years of construction, one year is given to achieve designed through-put.

The calculations for the cash flow analysis are given in Appendix D and the results are

discussed in this section. Figure 5.1 is the graph of cumulative cash position.



Figure 5.1: Cash flow diagram

From this analysis the different representations of the plant profitability is calculated, these

are ROI, PBP, NPV, and IRR with their different definitions in Section 5.1.1. These results

are given in Table 5.4.

Table 5.4: Profitability results from cash flow analysis

Profitability representation Value

ROI (return on Investment) 60.5%

PBP (payback period) 3.98

NPV (net present value) R 2.36 billion

IRR (internal rate of return) 352%

-2

0

2

4

6

8

10

12

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22

Cas

h flo

w (R

bill

ions

)

Years (yr)Cumulative cash flow Cumulative discounted cash flow



According to Peters et al. (2004) the minimum rate of return (mar) for projects that increases

the capacity of an existing plant and has an established corporate market position is

between 8 and 16% per year. For this project the ROI is calculated as 60.5% which is

significantly higher than the minimum rate of return which provides a favorable investment to

the investor.

For the payback period to be acceptable it should be less than or equal to the value

calculated by Equation 5.4. The value calculated from this equation is 3.95 years which is

slightly lower than the calculated PBP from the cash flow analysis. A positive NPV is

obtained from the cash flow analysis which is an indication of the money earned from the

project regardless of the investment earnings. Finally, as seen in Table 5.4, the NPV is

greater than zero which means the IRR should be calculated to give an indication of the

discounted cash flow rate of return.

5.4. Economic sensitivity analysis

The economic sensitivity analysis is done to investigate the effect of various factors on the

profitability indicators. The factors under discussion are:

• Revenue from a unit product.

• Inflation rate.

• Fixed cost.

• Variable cost.

• Tax rate.

These factors are varied by increasing and decreasing the factors with increments of 5% up

to 20% to obtain the effect on the NPV value. The sensitivity analysis is given in the graph in

Figure 5.2 with the calculations in Appendix D.



Figure 5.2: Economic sensitivity analysis

As seen from Figure 5.2, the factor that positively affects the NPV the most is the revenue

received from a unit ADU with an increase of 30.9% in the NPV as the revenue increase with

20%. The other factor positively affecting the NPV is the inflation rate. The factors that

negatively affects the NPV the most is the variable cost per ton product. This means that as

the raw material costs, utility costs, operating labour costs, ect. increase or decrease the

NPV will be affected the most. The other factors negatively affecting the NPV are the tax

rate and fixed cost per annum.

-100%

-80%

-60%

-40%

-20%

0%

20%

40%

60%

80%

100%

-20% -15% -10% -5% 0% 5% 10% 15% 20%

Pers

enta

ge c

hang

e in

NVP

Percentage change in variableRevenue from a unit product Inflation Fixed Cost Variable Cost Tax Rate



5.5. Recommendation for profitability

Although this project already shows a desirable profit, there is always room for improvement

without jeopardizing the quality of the product and safety on the plant. In this section some

recommendation are given to improve profitability. These recommendations are:

• Optimization of process:

After the construction of the additional equipment and the desired through-put is

achieved, it is necessary to optimise the process. This optimization will be done to

better understand the process conditions which can be improved to achieve optimum

efficiency of the process.

• Optimization of reagents:

At the completion of the process optimization, the reagent use should be optimised to

achieve a maximum amount of product with the least amount of raw materials.

Recycle streams should be considered to achieve this optimization.

• Minimum loss of solvent:

One of the most expensive reagents is the solvent used in the solvent extraction

section, thus measures should be considered to minimize the loss of this reagent.

• Optimal use of existing structures:

Since this project is an upgrade of an existing plant, it should be considered to use

existing processing equipment and administrative buildings on the plant instead of

replacing them. The existing equipment that is not used should be sold, if possible.

• Commission a processing plant for ADU:

At present the ADU is sent to NUFCOR for processing. It is recommended that an

economic feasibility study is conducted to compare the commissioning of a new ADU

processing plant to the outsourcing of this processing to NUFCOR.


143 Chapter 6: Safety and environment

Chapter 6: Safety and environment

From an economical point of view it is better to prevent accidents rather than to repair the

damage after the accident. This is why, when considering safety and environment, the focus

should be on preventing the hazards associated with the chemicals used, process

equipment and overall safety on the plant. In this chapter the importance of ensuring the

safety of the employees, the general public and surrounding environment are emphasized.

The sections under discussion in this chapter are:

• The overall safety specifications that should be considered for the plant and the

corrective measures taken in case of hazard occurrence.

• Hazard and Operatibility (HAZOP) level 1 study which ensures that the hazards

associated with the chemicals involved in the plant is understood properly.

• Environmental impact and management. This is done to ensure that the wastes

generated from the plant are treated in the correct manner to minimize the plants

impact on the surrounding environment.

• The plant location, which is important to determine the environment surrounding the

plant. A preliminary plant layout is discussed to ensure safety of the employees,

general public surrounding the plant, and the environment.

In all the sections the necessary occupational health and safety laws and standards should

be implemented.

6.1. Overall safety specifications

This section considers the overall safety specifications of the plant, which entails a

discussion about the possible hazards of the chemicals in each processing section as well

as some of the effects of these chemicals on the equipment used. The complete hazard

assessment of each chemical is done in Section 6.2 and only the most important hazards

are discussed in this section. There are preventative measures that can help reduce the

possibility of a hazard becoming an accident.



6.1.1. Leaching

In the acid leaching section the raw materials used are sulphuric acid, nitric acid, and

potable water. Both the sulphuric and nitric acid is extremely hazardous and corrosive

materials and human contact should be avoided at all cost. The correct signage should be

provided where these materials are used and stored as well as the correct personal

protective equipment (PPE) should be worn at all times when near the chemicals. The

sulphuric acid is diluted with the potable water, thus the correct procedure should be

followed when mixing these two chemicals together. The acid is added to the water, and not

the other way around, otherwise an extremely exothermic reaction will occur and could

cause an explosion.

The leaching equipment used is steam-heated, air-agitated, open pachucas tanks. The

hazards present for this equipment type is the heating steam and the fact that these tanks

are open with a large amount of hazardous chemicals inside. To prevent personnel falling

into the pachuca tanks the correct protective railings should be provided at the top of these

tanks. To ensure safe working conditions at this plant section the tanks should never be

operated at full capacity to prevent overflow of the tanks. The steam lines to the tanks

should be provided with the correct signage to indicate that the lines are hot and the lines

should be inspected regularly to prevent leakages.

The counter-current decantation process is also part of the leaching section. The magnafloc

used in this process is extremely slippery and spillage of this compound should be avoided.

In case of spillage, safety shoes with adequate grip should be worn and workers should

avoid the contaminated area. The processing equipment used is open thickeners. The

preventative measures for the thickeners are the same as that for the pachuca tanks, the

correct railings should be provided to prevent personnel from falling in and the correct

signage should be provided to indicate that the thickeners contain hazardous chemicals.



6.1.2. Ion exchange

The second unit in the process is the ion exchange unit where the uranium is loaded onto

the resin. The resin used is extremely slippery and spillage of this resin should be avoided.

In case of spillage, safety shoes with the adequate grip should be worn and workers should

avoid the contaminated area. The uranium complexes are a heavy metal and spillage of

these complexes will be dangerous to workers and the environment. Sulphuric acid is used

as eluant and is harmful to the workers. The acid and sulphate complexes that are present

in the system are corrosive which will damage the equipment..

The ion exchange unit consists of vertical pressure vessels with dished ends and safety

railings should be provided to the walkway at the top of the vessels to prevent workers from

falling down. To protect the environment and workers against spillage, containment walls

should be built around the units. The appropriate signage should be used to warn workers

against the corrosive nature of the chemicals in case of spillage. Special PPE should be

worn at all times in case of spillages or leaks.

6.1.3. Solvent extraction

The next processing step is the solvent extraction process which contains a number of

hazardous chemicals. The chemicals present in this section are the organic solvent and the

aqueous feed from the ion exchange section. The solvent is a composition of kerosene,

alamine® 336 (alkyl amine), and isodecanol. All these organic chemicals pose a fire hazard

that should be prevented at all cost. The solvent extraction section of the plant should be

isolated from the rest of the plant to prevent fire damage to the rest of the plant in case a fire

start in the solvent extraction section. These chemicals should be stored away from ignition

sources and the correct fire fighting equipment should be present in case of fire. The fire

fighting procedures are discussed in section 6.1.5. Exposure of these chemicals to

personnel and the environment should be kept to a minimum.

The processing equipment used in the solvent extraction section is mixer-settlers. The

potential hazards associated with this type of equipment are static and operational sparks

which act as ignition sources and should be strictly controlled. Sparks are prevented by

thoroughly grounding the equipment and to isolate the mechanical and electrical equipment

to enclose the sparks generated. Additional precautions are discussed in section 6.1.5.



6.1.4. Precipitation

Precipitation is the final unit of the process where ammonia and compressed air are used to

precipitate the product, ADU, out from the OK liquor. The OK liquor, coming from the

solvent extraction unit, and the ADU, is both radioactive and the ADU is corrosive. The

ammonia is toxic when inhaled, having a toxic affect on the respiratory system and other

internal organs. When working at the precipitation unit it is necessary to wear respiratory

equipment and safety goggles to protect workers against ammonia inhalation and ADU dust.

To minimise the environmental impact of these substances, all the equipment of the

precipitation unit is inside a building and access to the building is limited. The centrifuges

that is used to wash the ADU product causes hazardous noise which means ear protection

must be used and the noise should be monitored to make sure it does not exceed the noise

limit given by the occupational health and safety act. The necessary protective hand railings

should be provided at the top of the thickener. As mentioned, the ADU product is radioactive

and extra caution must be taken to ensure that workers are not exposed to the product be

means of leakage. Radiation control is discussed in detail in section 6.1.9. The equipment

must be checked regularly to prevent spills and leakage.

6.1.5. Fire fighting

One of the most dangerous hazards on the metallurgical plant is the possibility of fire. The

largest fire hazard is located at the solvent extraction section of the plant. Thus fire

prevention and control is crucial to the design of the plant. Fire prevention and control

includes the detection, suppression, and mitigation of existing fires.

The standard fire prevention and control systems include the following (BHP billiton, 2009:

673):



• Fire protection requirements in the design of plant and equipment.

• Training and induction for all personnel.

• Specialist training for high-risk areas.

• Dedicated response personnel.

• Centralised fire detection.

• Fixed fire suppression systems.

• Planned workplace inspections.

• Hazardous materials management.

Since there may already be existing fire prevention and control measures implemented at

the South Uranium Plant, it is necessary to access these measures and equipment with the

criteria listed above. If the requirements are not met the procedure and equipment should

be upgraded and new measures should be implemented for the upgraded section of the

plant.

Special attention is dedicated to the fire prevention and control on the solvent extraction

section of the plant, since the largest fire hazard is located in this section. The main cause

of fire that should be considered is static discharge. Static discharge can be avoided by

applying certain rules including (Watson, 2009):

• No flow velocities above 1 m/sec.

• Adequate grounding of all equipment.

• Use of PTFE diaphragm valves.

• Fibre flooring and walking grid.

• Pipes should be operated at full capacity to avoid friction between liquid and vapour.

Additional precautions that should be implemented for safety in the solvent extraction section

are a controlled entrance, specialized automated fire protection system, and a 15 m dead

zone around the section. The purpose of the controlled entrance is to prohibit personnel and

guests from entrance the section with equipment or accessories which may cause ignition.

All equipment and accessories are placed in a safety box at the entrance of the section and

this is controlled by a security guard.



The specialized automated fire protection system consists of both detection and

extinguishing equipment. Each vessel containing solvent is equipped with three fire

detection probes. A fire will only be registered if two of these probes detects a fire to prevent

accidental release of extinguishing foam which causes total loss of solvent. Each tank is

equipped with three extinguishing foam release points to ensure adequate dispersion of

extinguishing foam.

The 15 m dead zone is to minimize the damage to surrounding process equipment in case of

fire. Additional purpose of this dead zone is to isolate all the electrical transmitters from the

hazardous environment. All the electrical signals are transferred to Zener barriers, which is

located outside the dead zone, to prevent electrical sparks inside the hazardous

environment. A Zener barrier is an intrinsically safe explosion-proof system for electronic

instrumentation equipment (Fuji electric systems Co., Ltd., 1989).

It must be kept in mind that the necessary fire fighting equipment and procedures should be

in place to comply with insurance guidelines to ensure the safety of the investment. All

these prevention and protection measures should be implemented to ensure a safe working

environment for personnel and protection of equipment in the solvent extraction section and

the entire plant.

6.1.6. Training and personal safety

AngloGold Ashanti management and personnel are driven by the company values. The first

value concerns safety and is given as (Cutifani, 2009):

“We place people first and correspondingly put the highest priority on safe and healthy

practices and systems of work. We are responsible for seeking out new and innovative ways

to ensure that our workplaces are free of occupational injury and illness. We live each day

for each other and use our collective commitment, talents, resources and systems to deliver

on our most important commitment ... to care.”

To achieve this important value the company should provide training for all personnel and

visitors before allowing them to enter the working environment. The training should include

the following categories:



• General process information training.

• Risk assessment training.

• General plant safety training which include PPE, radiation, fire, and gas leak training.

• Plant specific training which includes chemical hazard training, and handling of tools

and equipment.

The general process information training is important to inform all personnel of the process

and layout of the entire plant. The overview of the process will provide the personnel with a

better understanding of potential hazards associated with the plant environment. The

personnel must know the layout of the plant to ensure that safety equipment and assembly

points are easily found. To further equip the personnel with the necessary skills to ensure

safety on the plant, the risk assessment training is provided.

The general plant safety training is important to protect the personnel and equipment on the

plant. No visitors are allowed on the plant without the proper plant safety training to ensure

public safety. This training entails PPE, radiation, fire, and gas leak training. The PPE that

should be worn at all times is; hardhat, safety goggles, safety boots, ear protection, gloves, a

safety overall, and respiratory protection at the ADU section. The radiation, fire and gas leak

training is required to ensure that safety equipment is used in the correct manner to protect

the personnel and equipment.

Each process unit should have their own plant specific training to inform the personnel

working in that section of the specific hazards involved and how to handle these situations.

In this training, a more detailed assessment of the chemical hazards and chemistry for each

section is given. Tools and equipment training is important to avoid operation errors and to

protect the specialised equipment used in the section.

6.1.7. Danger zones and signs

The most efficient safety measure is to use safety signs to indicate the potential hazards at

the relevant location. If there are dangerous zones present on the plant that is a high risk to

personnel, it should be clearly indicated using the appropriate signage. This signage should

provide information regarding the nature of the hazard as well as the safe operating

procedures when working in this dangerous zone. It is also very useful to provide safety

signs which indicate the necessary safety measures such as the PPE that have to be worn

and the location of emergency assembly points.



All safety equipment, such as fire extinguishers and breathing apparatus, must be indicated

with the correct signage. It is essential to make sure all personnel understand, are aware of

and obey the signs used on the plant, by means of general and plant specific training and

continuous inspection. Some examples of the necessary signage are shown in Figure 6.1.

Figure 6.1: Examples of necessary signage.

6.1.8. Emergency response plan

Although more resources are dedicated to the prevention of incidents, it is also important to

be prepared when accidents do happen. The emergency response plan is provided to

minimise the damage to equipment and injuries to personnel that might be caused by

accidents. The emergency response plan should include a response plan to ensure the

safety of personnel, a response plan to minimise damage to equipment and to control the

accident, and a response plan to notify the surrounding residential and business areas that

might be affected.

The safety of personnel is ensured by using a variety of alarms to notify all personnel of the

accident or potential hazard. Different alarms should be used for different accidents. When

the personnel are notified, everyone must follow the evacuation procedure provided and

gather at the assigned emergency assembly points. At these assembly points roll call must

be taken to ensure all personnel are safe. Personal safety equipment, such as breathing

apparatus, is provided on the plant for additional protection if the assembly points can not be

safely reached.



To protect the equipment on the plant in case of an accident and to contain the hazardous

chemical spills, emergency shutdown procedures should be implemented. Each section on

the plant should have a specific shutdown procedure which should be integrated into the

entire process. The hazards of chemical spills can further be contained by building

containment walls around the equipment with a high potential of chemical spill.

If the accident can not be controlled, the necessary safety services should be notified to

respond as quickly as possible. To ensure quick communication to the safety services,

dedicated phone lines should be provided on site.

6.1.9 Radiation control Radiation is an unseen hazard and little information is available to the public, therefore this

hazard is often misunderstood which means more resources should be dedicated to the

control of radiation exposure. The International Commission on Radiology Protection (ICRP)

provides a standard dosage limitation measured in sieverts. The ICRP states that the

radiation limit for workers is 20 mSv per annum while the limitation for the public is 1 mSv

per annum (BHP Billion, 2009).

The radiation limits of each employee should be monitored each month by means of a

continuous personal dosimeter which monitors radon gas. At the end of each month a report

should be compiled, listing all the radiation exposure results, and sent to the National

Nuclear Regulator (NNR) (AngloGold Ashanti, 2007:41). Another way of determining the

dosage that the workers have been exposed to, is defined by the ICRP and involves the

determination of the different ways in which workers might be exposed to radiation. The

radiation levels must be measured in the identified areas and together with the time spent in

the exposed areas, the dosage can be determined using an internationally accepted

conversion factor (BHP Billion, 2009). In Table 6.1 some examples of radiation dosages are

given (NECSA, 2009).



Table 6.1: Examples of radiation dosages

Dosage Effect

0.01 mSv Radiation dosage received by patient having his/her teeth X-rayed.

0.01 mSv Radiation dosage received by patient having his/her lungs X-rayed.

2 mSv Annual dose of cosmic radiation received by a person working in an

aeroplane.

4 mSv Average annual radiation dose for South Africans caused by indoor radon, X-

ray examinations, etc.

100 mSv Highest permitted dose for a radiation worker over a period of five years.

1 000 mSv The dose which may cause symptoms of a radiation sickness if received

within 24 hours.

6 000 mSv The dose which may lead to death when received all at once.

To ensure that the workers do not reach the exposure limit, efficient ventilation systems

should be installed, and old buildings which might cause leaks should be monitored and

repaired. Internal radiation limits should be established, which is lower that the standard

limit, and if this limit is reached the workers should be moved to another working area where

the exposure risk is lower (AngloGold Ashanti, 2007:41).

6.2. HAZOP level 1

Hazard and Operability Studies (HAZOP’S) is an integral part of the design process of a new

plant or piece of equipment. These studies analyse the potential hazards and operational

problems which is divided into six sub-sections. The most important sub-sections for the

preliminary design of a new plant are:

• HAZOP level 1: Examination of the entire input-output structure to access the health

and safety considerations of the raw materials used and products produced.

• HAZOP level 2: A study which investigates the possible dangers associated with

specific plant equipment. This will ensure adequate protective measures and

environmental safety.

• HAZOP level 3: Any risks associated with the general operation of the process are

examined and necessary control measures are applied.

The three HAZOP levels above should be carried out in a certain sequence; this sequence is

displayed in Figure 6.2.



Figure 6.2: HAZOP structure for plant

For the scope of this conceptual design, only a HAZOP level 1 study is completed in detail.

A HAZOP level 1 study is important to understand the overall operation of the process and

the hazards associated with the raw materials and products. It is important to complete this

study at an early stage of the design process, since it is more economic favourable to

prevent hazards than it is to repair the damage after an accident. The HAZOP level 1 study

is divided into separate sections, these sections are discussed below.

6.2.1. Project definition

This project entails the upgrade of the current South Uranium Plant (SUP) of AngloGold

Ashanti. The ore processing rate needs to be increased from 240 000 to 360 000 ton/month.

To accommodate the additional feed the equipment should be upgraded or redesigned. The

process units under discussion are leaching, ion exchange, solvent extraction and

precipitation. The operating time is 8150 hours per year, which allows for maintenance time

and unscheduled shut-downs. Table 6.2 gives the projected production rate of the ADU and

gold for the upgraded plant.

Table 6.2: Projected production rate for the upgraded plant

Product Production rate (ton/annum)

ADU 1068

Gold 2.16

This increase in production, although expensive, is expected to provide long-term

economical benefits.



6.2.2. Process description

A detailed description of the process is provided in Chapter 3 hence a very short description

of the process is given in this section.

Unit 1 (U01): Leaching and CCD

The uranium extraction process starts with the leaching of uranium containing ores. The

uranium is liberated using nitric acid as oxidant and sulphuric acid as lixiviant. The leach

product stream contains both solids and liquids which are separated using counter-current

decantation. Magnafloc 90L is used as a flocculant to enhance the efficiency of

sedimentation.

Unit 2 (U02): Ion exchange

The liquid product from the counter-current decantation unit contains the uranium product

and other impurities. The uranyl sulphate ions are selectively adsorbed onto the resin,

Ambersep TM400. Using eluant, the uranyl sulphate ions are desorbed and sent to the

solvent extraction unit. The desorbtion process results in an increased uranyl sulphate ion

concentration and the removal of some impurities.

Unit 3 (U03): Solvent extraction

The solvent extraction process consists of four sections which includes extraction,

scrubbing, stripping and regeneration. The extraction section removes the uranyl sulphate

complexes from the aqueous phase into the organic phase. Most of the iron sulphate

complexes are entrained into the organic phase during this section. The scrubbing section

removes the entrained iron sulphate complexes from the organic phase. Stripping, using

ammonium sulphate, of the uranyl sulphate complexes from the organic phase into the

aqueous phase takes place in the next section. The organic solvent is regenerated in the

final section from where it is recycled to the first section.



Unit 4 (U04): Precipitation The final unit of the uranium extraction process is the precipitation of the product, ammonium

diuranate (ADU). The OK-liquor is sent from the solvent extraction unit to the precipitation

tanks where ammonia gas is combined with compressed air, and added to precipitate the

ADU. The pH has to be controlled at 7.5 to maintain optimum precipitation. The precipitated

ADU is washed in centrifuges and the final project is stored in ADU storage.

6.2.3. Assessment of chemical hazards

The assessment of the chemical hazards only focus on individual hazards associated with

the raw materials and the products. The services, instrumentation, and plant equipment is

not considered within this assessment. The Chemical Hazard Proforma HS1A serves as a

guideline in this assessment. Table 6.3 is the chemical guide used for the HS1A performa,

Table 6.4.

Table 6.3: List of chemicals

Chemical compound Physical state

Quantity (Inventory or through-put) (ton/annum)

A ADU Solid 624

B Alamine® 336 (alkyl amine) Liquid 9.26

C Ammonia Gas 174.67

D Calcium oxide Solid 83 016.98

E Ferrous sulphate Liquid 3 548.11

F Isodecanol Liquid 9.26

G Kerosene Liquid 283.28 m3

H Magnafloc 90L Liquid 215.98

I Nitric acid Liquid 13 844.41

J Nitric oxide Gas 2 149.38

K Potable water Liquid 3586

L Resin (Ambersep 400) Solid 123.25

M Silicon dioxide Solid 501

N Sodium carbonate Solid 84.16

O Sodium hydroxide Solid 168.32

P Sulphuric acid Liquid 143 005.73

Q Uraninite Solid 0.25



Table 6.4: Chemical hazard data sheet (HS1A)

Hazard potential ‘-‘ Key: ‘K’ ‘No.’

Insignificant hazard Hazards known and understood See numbered notes

A B C D E F G H I J K L M N O P Q

Explosion and flammability hazards

Fire K K - - - K K - - - - - - - - - -

Deflagration/Detonation - K K - - - K - K - - - - - K K -

Electrical static - K - - - K K - - - - - - - - - -

Reactivity/stability hazards - - K - - K - K - - - - - - - - -

Immediate health hazards

Inhalation toxicity K K K K K K K - K K - - K K K K K

Skin absorption K K - K K K K - K K - - K K K K K

Corrosive - - K K - - - K K K - - - - K K -

Chronic health hazards

Digestive K K - K K K K - K - - K K K K K K

Sensitiser K K - - - K K - K K - - - - K K -

Continual K - - - K - K - K K - - K K K K K

Other health hazards

Odor - K K - - K - - K K - - - - - - -

Radiation K - K - - - - - - - - - - - - - K

Environmental hazards

Aqueous K K K - - K - - K K - - - - K K -

Gaseous - - K K - - - - - K - - - - - - -

Ground K K - K - K - - K K - K - - K K -

Hazard breakdown products K K - - K K - - K - - - - - K K K

6.2.4. Assessment of chemical interactions

After the individual hazards have been assessed, the chemical interactions between the

chemicals and construction materials should be considered to prevent any unknown and

undesirable reactions within the process. This assessment must be taken into consideration

when designing the entire plant. The Chemical Interaction Proforma HS1B, Table 6.5, is

used with the chemical guide, Table 6.3, to evaluate these interactions.



Table 6.5: Chemical interactions data sheet (HS1B)

Hazard potential ‘-‘ Key: ‘K’ ‘No.’

Insignificant hazard Hazards known and understood See numbered notes

Chemical or group of chemicals

B C D E F G H I J K L M N O P Q

A ADU A - - - - - K - K K - K - - - K -

B Alamine® 336 B - K - - - - K K - - - - - - -

C Ammonia C - - - K - K - - - - K K - -

D Calcium oxide D - K K K K - K - - - - K -

E Ferrous sulphate E - - - - - - K - - K - -

F Isodecanol F - - K K - - - - - - -

G Kerosene G - K K - - - - - K -

H Magnafloc 90L H - - - - - - - - -

I Nitric acid I K - K K K K K K

J Nitric oxide J - K - - K - K

K Potable water K - - - - - -

L Resin L K - K K -

M Silicon dioxide M - - - -

N Sodium carbonate N - K -

O Sodium hydroxide O K K

P Sulphuric acid P K

Q Uraninite Q

Chemical compound

A B C D E F G H I J K L M N O P Q Material of construction

Carbon steel - - K K - - - K K K - - - - K K -

Concrete - - K - - - - - K K - - - - K K -

Glass reinforced

plastics - - - - - - - - - - - - - - - - -

Rubber - - - - - - - - - - - - - - - - -

Stainless steel316L - - - - - - - - - - - - - - - - -

Vacuumed infused

fibreglass - - - - - - - - - - - - - - - - -



6.2.5. Means of handling

When handling hazardous chemicals the possibility of an accident is inevitable if the hazards

are not carefully evaluated. The means of handling Proforma HS1C is used to evaluate the

handling hazards of the chemicals on the plant. The HS1C for is given in Table 6.6.

Table 6.6: Means of handling (HS1C)

Chemical or chemical compound

A B C D E F G H I J K L M N O P Q

Storage K - K K K K K K K K - - K K K K K

Transport K K K K - - K K K K - - - - K K K

Materials handling K - K K K K K K K K - - K K K K K

Extreme process conditions K - - K K - K - K K K - K K K K K

Other construction materials - - K K - - - - K K - - - - K K -

Decontamination K - - K - K K K - K - - K K K - K

Gaseous emissions - - K K - - K - K K - - - - - - -

Liquid release K K - - K K K K K - - - - - K K -

Disposal wastes K K K K K K K - K K - K K K K K K

Destruction of material K - - - - - - - - - - - - - - - K

Quality control K K K - - K - - K K - - - - - K -

Emergency procedures K K K K K K K K K K K K K K K K K

Plant layout, spacing, access K K K K - K K - K K - - - - K K -

Area classification K K K K K K K K K K K - K K K K K

Provision of services K - K - - - K - K - - - - - - K -

Codes of practice K K K K K K K K K K K K K K K K K

As seen in Table 6.6 there is an abundance of hazards associated with the handling of the

chemicals on the South Uranium Plant. This should be taken into careful consideration

when designing and upgrading the plant.



6.3. Environmental impact and management

According to the Environmental Conservation Act of 1989, the protection and rehabilitation of

the natural environment should be enforced by all. To accomplish this, the environmental

statement of the HAZOP level 1 study is done. This environmental statement will include a

waste block diagram analysis, handling and disposal of wastes, accidental releases of

hazardous materials and other occupational hazards to the environment.

6.3.1. Waste block diagram analysis

A waste block diagram for each processing unit is given in this section to make sure all raw

materials, wastes and products are accounted for. A typical waste block diagram is given in

Figure 6.3.

Figure 6.3: General waste block diagram structure

The first processing step is leaching where the valuable minerals are dissolved using lixiviant

and oxidizing agent. Figure 6.4 is an illustration of the waste block diagram for the leaching

process.



Figure 6.4: Leaching waste block diagram

The solid waste is transported to the gold processing plant next to the SUP, while the

pregnant leach liquor is sent to the ion exchange unit. In the ion exchange unit the

concentration of U3O8 is upgraded using a resin. Figure 6.5 shows the waste block diagram

for the ion exchange unit.

Figure 6.5: Ion exchange waste block diagram

The barren liquor is recycled to the CCD unit in the leaching section to provide wash solution

for the CCD process. The eluate is sent to the solvent extraction section where the

concentration of U3O8 will be further upgraded. The waste block diagram of the solvent

extraction section is shown in Figure 6.6.



Figure 6.6: Solvent extraction waste block diagram

The unwanted ADU gunk is added back into the process at the leaching section to be

processed again. The FeSO4 waste should be disposed of in the correct manner to ensure

environmental safety. The OK liquor proceeds to the precipitation section. Figure 6.7 is a

representation of the precipitation waste block diagram.

Figure 6.7: Precipitation waste block diagram

The SO42- rich waste is recycled to the stripping unit in the solvent extraction section to strip

the U3O8 from the solvent. The water waste contains harmful chemical and should be

disposed of in the correct manner.



6.3.2. Handling and disposal of wastes

The greatest concern when handling the waste of a uranium plant is the radioactive nature of

the materials that are released, in solid form, into the surrounding air and into the water

environment. It is important to consider all forms of waste which include solid, liquid and air

borne waste. When planning how to handle and dispose all waste it is essential to consult

the various safety standards and principles that should be met. International atomic energy

agency (IAEA, 1995) provides the following principles for the management of radioactive

waste:

• Protection of human health.

• Protection of the environment.

• Protection beyond national borders.

• Protection of future generations.

• Burdens of future generations.

• National legal framework.

• Control of radioactive waste generation.

• Radioactive waste generation and interdependencies.

• Safety of facilities.

To further assist with the design of the waste disposal units, the different wastes can be

classified in to four levels, shown in Table 6.7 (NECSA, 2009).

Table 6.7: Waste classification

Classification Description

Very low-level waste (VLLW) Contains very low concentrations of radioactivity,

originating from the operation and decommissioning of

nuclear facilities.

Low- and intermediate level

waste (LILW)

Contains concentrations or quantities of radionuclides

above the clearance levels established by the regulator.

High-level waste (HLW) Contains heat-generating radionecluides with long- and

short-lived radionuclide concentrations.

NORM Contains low concentrations of naturally occurring

radioactive materials.



The waste generated on this plant is classified as NORM waste. The South African

department of minerals and energy (DME) provides certain regulations for the waste

management of radioactive waste and there are an additional three policies that should be

considered. These policies are:

• The Radioactive waste management policy and strategy for the Republic of South

Africa (2005).

• The Principles of Radioactive Waste Management (IAEA, 1995).

• Management of Radioactive Waste from the Mining and Milling of Ores (IAEA, 2002).

• Storage of Radioactive Waste, Safety Guide No. WS-G-6.1 (IAEA, 2006).

A few of the considerations that should be addressed are to ensure an adequate ventilation

system and that equipment is enclosed to prevent wind dispersion of contaminated

materials. The unit where the product is formed should be located in a separate building

with limited access to minimize the exposure of workers to radiation.

On the existing plant a waste management plan is already in practice. This system will be

used for the already processed feed and the upgraded feed. An upgrade might be

necessary when more information is available. The gold slurry produced at the CCD section

should be neutralization to increase the pH from 1.8 to 10.5, lime slaking is used for this

neutralization and this is done in one of the existing pachucas.

The upgrade of the plant involves a new process using nitric acid as oxidizing agent, causing

the release of nitrate ions. Although this waste is less hazardous than the waste that is

already produced, it can not be ignored. A process is already developed which uses

biochemical operations to remove nitrogen from the wastewater. This process is known as

denitrification and involves the reduction of nitrate-N to N2 by heterotrophic bacteria that

grow in the absence of oxygen and the presence of nitrates. The nitrates act as the terminal

electron acceptor and the micro-organisms used grow in suspension in the liquid that is

treated. Because the wastewater treatment bioreactors is depended on the way in which the

micro-organisms grow, for denitrification a continuous stirred tank reactor (CSTR) is needed

(Grady, 1999:216). In the USA 100 wastewater treatment plants use methanol in the

denitrification process which serves as a carbon source for bacteria (Methanol Institute,

2009). These wastewater treatment facilities serve as an indication that this method of

wastewater treatment is efficient and can be used.



6.3.3. Accidental releases of hazardous materials

It is not possible to completely avoid accidental releases of hazardous chemicals. Therefore

it is necessary to have adequate preventive measures and emergency response in case of a

release. Table 6.8 lists the hazards materials that may lead to potential spillage.

Table 6.8: Accidental releases

Materials Impact Avoidance Measures or Actions

ADU Exposure to radiation Signage and confined work area.

Acids Chemical burn Signage and containment dams around

equipment.

Magnafloc 90L Slippery surfaces Signage and correct PPE.

Kerosene Fire hazard Signage and security measures must

be taken.

Ore Dust dispersion and small dose

radiation

Signage, correct PPE and enclosed

structures.

Strict implementation of the avoidance measurements should be enforced to ensure a safe

working environment for personnel and the natural environment.

6.3.4. Rehabilitation and decommissioning of plant

The decommissioning of a plant is required by the government in order to protect the

environment and reduce the exposure to radioactive materials. The decommissioning of a

plant starts after plant operations seized and involves the decontamination and dismantling

of equipment and the disposable of all radioactive materials. Decommissioning can be

described as the administrative and technical actions that should be taken to remove the

facilities which have been exposed to radioactive material on such a scale that it can be

considered a safety hazard. Decommissioning of facilities is becoming a major issue and

such be considered in every plant design (IAEA safety standards, 2006: 1).



The decommissioning of a plant can be divided into three phases. The first phase involves

the plant shutdown, followed by the disposal of process materials. In this first phase the

plant site is made safe and prepared for care and maintenance. During the second phase

some of the process equipment is dismantled, component sizes are reduced and the

equipment is decontaminated. In the final phase the buildings on site is decontaminated for

de-regulations (NECSA, 2009).

It is recommended that a preliminary decommissioning plan should be submitted to the

government five years before the projected end of operations. This plan should include the

cost estimation, discussion of the major technical actions that will need to be taken and

lifetime of the decommissioning project (Fentiman et al, 2009:1).

Some guidelines are provided by NECSA (2009) that can be followed when the preliminary

decommissioning plan is constructed.

• A complete inventory of all facilities, equipment and materials that is exposed to

nuclear radiation should be complied.

• The materials and processes that need to be decontaminated and dismantled should

be categorised. These categories are defined by the extent to which the materials

and equipment should be processed.

• The capacities and cost of each of the processes are determined.

• The nuclear exposure of all the equipment and materials should be assessed with

specially developed software packages.

6.4. Plant layout and positioning

The plant layout and positioning is done to ensure that enough land is available for the new

upgraded equipment and to get an overall representation of the positioning of all the different

units. It is important to note that the equipment sizes are not on scale but the plant area is a

more accurate representation. The layout was constructed using the air photo obtained from

the program Google Earth and the old plant layout received from William Manana which is

included in Appendix E. The new plant layout is given in Figure 6.8 and the legend included

assists in the interpretation of the figure.



Figure 6.8: Plant layout


167 Chapter 7: Process control

Chapter 7: Process control

In the previous sections the process is design assuming steady-state operating conditions using

a specified feed. It should however be kept in mind that there are multiple variations in the

process conditions due to external disturbances on the process. These variations influence the

process which results in dynamic operating conditions. Since the process is designed from

steady-state assumptions it is important to manipulate the controllable variations to ensure that

the process conditions are within the design bounds.

When the process conditions are not within the design bounds, the process will operate under

either hazardous conditions or production quality loss conditions or both. These conditions

should be avoided since it will endanger the environment and plant personnel, damage process

equipment and result in economical loss. For these reasons it is important to have an effective

integrated control system to keep the process conditions within the design bounds.

The following aspects are addressed in this chapter:

• An integrated control strategy for the entire plant.

• A detailed process control for the solvent extraction section.

• Specific safety considerations in the control of the solvent extraction section.

When all these aspects are addressed, the basic control for the uranium extraction plant is

described. These control strategies will be used to optimise the plant operations and ultimately

achieve optimum plant operating conditions.

7.1. Plant wide control

The plant wide control section will discuss the control objectives and strategies for each unit in

the process. A control schematic is shown for each of the important equipment used in each of

the units



7.1.1. Unit 1: Leaching, CCD and neutralization

In this section the control strategies for the leaching, counter-current decantation and

neutralization processes are given and discussed.

Leaching Firstly the control for the leaching section is discussed. In the leaching section the uranium is

liberated through several chemical reactions. To ensure the effective operation of this section

all parameters influencing reaction kinetics must be controlled. The following control objectives

for the leaching section are identified:

• Reactant flow rates

• Pachuca pH

• Pachuca temperature

• Level of agitation

In order to control these process conditions, control strategies are devised. These control

strategies are shown in the control schematic in Figure 7.1.



Figure 7.1: Control schematic of the leaching section

In Figure 7.1, four control loops are seen, each one controlling one of the identified control

objectives. Figure 7.1 shows three reactant streams fed into the first pachuca; the ore feed,

H2SO4 feed and the HNO3 feed. Of these three reactant feeds only the flows of the H2SO4 and

HNO3 feed can be controlled since a constant ore feed in the slurry form is received from the

mills of the gold extraction plant. The H2SO4 feed provide the H3O+ ions required in the process

and is therefore used to control the pH of the reactor system which is discussed later.



It is decided to control the flow for the HNO3 feed stream as a constant ratio in regards to the

received ore feed flow. This is achieved by implementing a ratio control loop between the ore

feed and HNO3 feed streams. The volumetric flow of the ore feed is measured (FI 01) and

transmitted to control unit (FRC 01) where the flow is multiplied with the required ratio to obtain

a set point for the HNO3 feed flow. The volumetric flow rate of the HNO3 feed is also

continuously measured and transmitted to the control unit (FRC 01), which will then control the

valve opening according to the actual flow and calculated set point. This control loop operates

as a direct controller which implies that the valve opening should increase when the volumetric

flow rate of the ore feed increases. Since the disturbance (ore feed volumetric flow) is

measured and accounted for before it enters the process (pachuca), the control loop is

feedforward control.

As mentioned the pachuca pH is controlled using the H2SO4 feed flow. The pH is measured at

the exit stream from the pachuca. This measured pH represents the pH of the whole pachuca

contents due the intense agitation and mixing. The pH value is transmitted to the pH control

unit (PHC 01) from which a set point for the H2SO4 volumetric flow rate is controlled. This

controlled set point and the measured actual volumetric flow of the H2SO4 feed (FI 03) is used to

control valve opening. This concept where the flow set point rather than the valve opening is

controlled is known as cascade control and is widely implemented for liquid systems (Svrcek et

al., 2006: 131-135). It should be noted that the pH control is implemented on the first two

pachucas only, since the most H3O+ is used in these initial stages. The overall control loop

operates as a direct controller, since the H2SO4 flow rate should increase when the pH

increases. This control loop is seen as a feedback control loop, since the disturbance is

measured downstream of the process.

The next important parameter to be controlled is the temperature of the pachuca content which

has a significant effect on the reaction kinetics. It is assumed in the design that a temperature

loss of 0.5 °C is experienced over each pachuca, from which it is derived that the temperature

should be controlled at each fourth pachuca. It is imperative to keep the temperature in the

pachuca above 30 °C to achieve the reaction rate the process is design for. Therefore the

lowest temperature in the pachuca (at the top) is measrured to control the temperature above

30 °C. This temperature measurement is transmitted to the temperature control unit (TIC 01)

which controls the steam feed to the pachuca by implementing a cascade control loop. This

control loop operates as a reverse controller since the steam flow should be reduced if the

temperature increases. This control loop is seen as a feedback control loop.



The pachuca contents is continuously agitated using air agitation which ensures adequate

mixing. The air entering at the bottom of the pachuca is controlled at or above a pressure of

6 bar to ensure adequate mixing. The pressure control is achieved by measuring the pressure

of the air where it enters the pachuca and transmitting it to the pressure control unit (PIC 01).

The pressure control unit (PIC 01) controls the valve opening to achieve the desired pressure.

The compressed air is supplied at 6 bar by the gold extraction plant and introduced to each of

the pachucas. This control loop operates as a reverse controller since the air flow should be

reduced if the pressure increases. This control loop is seen as a feedback control loop.

Counter-current decantation

In the following section the control of the counter-current decantation section is discussed. In

this section the leach product is washed to effectively remove the dissolved uranium in the liquid

phase from the solids. This is achieved using a series of thickeners which results in solid-liquid

separation. This process requires long residence times in thickeners with large capacity. The

most important parameters in controlling the thickeners are:

• Residence time

• Wash ratio

• Solid content of pregnant leach liquor

• Rake resistance

Each of these control objectives are controlled using a practical control strategy. These control

strategies are shown in a control schematic in Figure 7.2.



Figure 7.2: Control schematic of the counter current decantation section

The control schematic of Figure 7.2 shows three feed streams, the leach product, wash solution

and floculant. The residence time is an important parameter and therefore the flow rates into

the thickeners must be kept as stable as possible. Further it is desired to achieve a low solid

content in the pregnant leach liquor to prevent fouling of the resin in the ion exchange column.

To ensure effective washing, the wash ratio is controlled at an optimum. The rakes gather the

slurry at the bottom of the thickener to the underflow, which at times can experience high

resistance due to the thickness of the settled slurry.

The residence time of the thickeners are fixed and determined by the feed flow rate of the leach

product and wash solution. Since a fixed wash solution to leach product ratio is used, only the

leach product feed is controlled to achieve the desired residence time. The leach product feed

flow is measured (FI 01) and transmitted to the flow controller (FIC 01) which manipulate the

valve opening to reach the determined flow. This control loop operates as a reverse controller

since the leach product flow should be reduced if the flow increases. This control loop is seen

as a feedforward control loop.



The wash ratio is controlled by manipulating the set point of the wash solution feed flow rate

according to the set point of the leach product feed rate. The wash ratio is designed at one,

which will give an efficiency of 99.99%. The set point of the wash solution flow rate is not

directly dependent on the leach product flow rate, therefore a process upset introduced by the

leach product flow rate is not duplicated in the wash solution. The flow rate for the wash

solution is measured (FI 02) and transmitted to the flow controller (FIC 02) which controls the

valve opening to manipulate the wash solution feed rate. This control loop operates as a

reverse controller since the wash solution flow should be reduced if the flow increases. This

control loop is seen as a feedforward control loop.

It is important to keep the solid content of the pregnant leach liquor below the prescribed

50 ppm. This is achieved by manipulating the floculant addition. The solid content of the

pregnant leach liquor is determined by sampling this stream once a shift for analysis. The

results of the analysis will determine the set point for the floculant addition. The set point of the

floculant addition will vary greatly in the initial commissioning stage, however this variable will be

optimised for the specific process and will eventually require less sampling. This control loop

operates as a direct controller since the floculant addition should be increased if the solid

content of the pregnant leach liquor increases. This control loop is seen as a feedback control

loop.

The thickness of the settled slurry may cause equipment damage since the rake is driven from

the centre of the thickener. This resistance is observed in the amps which the electric motor

requires. This effect is used to control the rake and reduce the probability of equipment

damage. The rake is controlled by hydraulically lifting the rake and allowing it to pass over most

of the settled slurry whenever high amps are measured. This control loop operates as a direct

controller since the rake height should be increased if the amps increase. This control loop is

seen as a feedback control loop.



Neutralization

The last process in this unit is the neutralization where the slurry product from the CCD section

is neutralised before it is sent to the gold extraction plant. This process is carried out in an

existing pachuca where the pH of the contents is increased to 10.5 by the addition of slaked

lime. The following control objectives are identified.

• Pachuca pH

• Level of agitation

The neutralised slurry is treated with cyanide at the gold extraction plant, which releases toxic

gases if the received slurry is acidic. Therefore it is important to control the neutralization

process effectively to ensure the safety of the environment. In Figure 7.3 a control schematic

for the neutralization process is shown.

Figure 7.3: Control schematic of the neutralization process



In Figure 7.3 the feed streams to the neutralization section is shown; the slurry feed, slaked lime

and air feed. The slurry feed is received from the CCD section and cannot be controlled due to

the immense influence it will have on the CCD steady-state operation. Adequate agitation is an

important factor to ensure efficient neutralization. The pH of the product must be 10.5 which is

acceptable for the gold extraction plant.

The pH of the pachuca content is controlled using the slaked lime feed flow rate. The pH of the

pachuca is measured (PH 01) and transmitted to the pH controller (PHC 01) which then controls

the flow rate of the slaked lime feed. The flow rate is controlled using a cascade control method

as discussed for Figure 7.1. This control loop operates as a reverse controller since the slaked

lime feed flow rate should be increased if the pH decreases. This control loop is seen as a

feedback control loop.

To ensure adequate agitation the air pressure of the air feed at the bottom of the pachuca is

controlled at 6 bar. This is achieved by measuring the air pressure (PI 01) and transmitting it to

the pressure controller (PIC 01) which controls the valve opening to control the pressure. This

control loop operates as a reverse controller since the valve opening should be increased if the

air pressure decreases. This control loop is seen as a feedforward control loop.

7.1.2. Unit 2: Ion exchange

The pregnant leach liquor from the CCD section is fed to the ion exchange system where the

concentration of uranyl sulphate complexes is increased. This is achieved by the selective

adsorption of these complexes onto resin, and after saturation of the resin the complexes are

desorbed during the elution stage. The liquid eluant (diluted H2SO4), used for the elution of the

uranyl sulphate complexes, has a lower volumetric flow rate to ensure a higher concentration.

This process is operated in a semi-continuous procedure in which each column of resin goes

through four separate stages; adsorption, back wash, elution and resin regeneration. The

control for each of these stages is discussed separately in the following sections. These

processes are commonly controlled according to time cycles which is calculated and determined

during the commissioning period.



When a column is in the adsorption stage, the uranyl sulphate complexes are adsorbed onto the

resin. The adsorption kinetics is fast and therefore an adsorption front is formed where the

adsorption takes place as described in Section 2.5.2. When this front reaches the bottom of the

column, break through is observed. The control objectives identified for this stage is given as:

• Pregnant leach liquor flow rate

• Column saturation

When the resin is saturated, the adsorption column is taken of the adsorption train to be back-

washed and eluted. The pregnant leach liquor flow rate influences the thickness of the

adsorption front. The control schematic for the adsorption stage is shown in Figure 7.4.

Figure 7.4: Control schematic of the adsorption stage of ion exchange

Figure 7.4 shows that the pregnant leach liquor is pumped into the resin-containing column and

exits to the next column in the train. The volumetric flow rate of pregnant leach liquor is

controlled to ensure adequate residence time and an acceptable adsorption front thickness.

The volumetric flow rate is measured (FI 01) and transmitted to the flow rate controller (FIC 01)

which controls the valve opening. This control loop operates as a reverse controller since the

valve opening should be decreased if the flow increases. This control loop is seen as a

feedforward control loop.



In many cases the required time to saturate the resin in a column is used as control mechanism

to indicate when the resin is saturated. Due to resin poisoning the required time should be

analysed frequently to reduce uranium loss. The saturation of the resin can however be

measured using its conductivity. In this method the solution temperature (TI 01), solution

conductivity (CI 01) and resin conductivity (CI 02) are measured to determine the level of

saturation of the resin (Serra & Solã, 1986:34). These measurements are transmitted to an

on/off controller (CIC 01) which determines the saturation of the resin and decides accordingly

to switch the pump on or off. This control loop operates as a multi-input-single output controller,

which is seen as a feedback control loop. When the resin is saturated the pump is switched off

and back-wash may start. The back-wash stage control schematic is shown in Figure 7.5.

Figure 7.5: Control schematic of the back-wash stage of ion exchange

During the back-wash stage the resin bed is fluidised to ensure the effective removal of small

solid impurities from the resin and remove pregnant leach liquor remaining in the resin bed.

Therefore the back-wash water flow rate is controlled to achieve fluidization. The flow rate is

measured (FI 01) and transmitted to the flow rate controller (FIC 01) which controls the valve

opening. This control loop operates as a reverse controller since the valve opening should be

decreased if the flow increases. This control loop is seen as a feedforward control loop. When

the back-wash stage is completed, the resin is eluted as shown in the control schematic in

Figure 7.6.



Figure 7.6: Control schematic of the elution stage of ion exchange

The eluant volume used to elute the resin is kept constant relative to the resin loading to

achieve the desired uranyl sulphate concentration. For this process, 11 bed volumes of eluant

is used with a residence time 15 minutes. To ensure this flow, the eluant flow rate is measured

(FI 01) and transmitted to the flow controller (FIC 01) which controls the valve opening. This

control loop operates as a reverse controller since the valve opening should be decreased if the

flow increases. This control loop is seen as a feedforward control loop.

The regeneration step of the resin is an important step to ensure a longer resin lifetime, and

better resin capacity. This stage is carried out in specific steps in a separate regeneration

column to prevent resin shock. Resin is pumped from the adsorption columns to this separate

adsorption column. Therefore the resin is first contacted with H2SO4, then water, then caustic

and backwards again. Figure 7.8 shows the regeneration control schematic.



Figure 7.7: Control schematic of the regeneration stage of ion exchange

Each of the flows through this regeneration column takes place in the different steps and is

controlled using flow controllers similar that seen in the elution and back-wash stages (Figure

7.6 and 7.7).

7.1.3. Unit 3: Solvent extraction

The eluate from the ion exchange unit contains an increased uanyl sulphate concentration and

is sent to an eluate feed storage tank with a large capacity. From the eluate feed storage tank,

the eluate is pumped to the solvent extraction unit where the concentration of the uranyl

sulphate complexes is increased further and impurities are removed.



This process is similar to ion exchange in that the uranyl sulphate complexes are transferred to

an organic phase and back to an aqueous phase in the different sections. The process consists

of the extraction, scrubbing and stripping stages with an extra regeneration section where the

organic solvent is regenerated. The main control objectives and strategies for the solvent

extraction unit are discussed below, while a more detailed discussion on the control of this unit

is given in Section 7.3. The following main control objectives are identified for the solvent

extraction unit.

• Volumetric flow rates of the various streams

• pH on the stripping stages

This process is designed to process a certain eluate feed flow rate with set ratios between the

flow rates of the aqueous and organic phases through the mixer-settler units. The stage

efficiency and phase separation in the stripping section is greatly dependent on the pH which

varies because of extraction reactions. A control schematic of the flow rates for the solvent

extraction unit is given in Figure 7.9.

Figure 7.8: Basic control schematic of the solvent extraction unit



Figure 7.8 shows that the eluate feed flow rate from the storage tank is controlled to ensure a

flow rate close to that for which the unit is designed for. From this the flow rate ratios in all the

sections are controlled to maintain the design conditions. The pH of each stage in the stripping

section is controlled using the caustic feed stream which is fed to each stage separately.

The eluate feed volumetric flow rate is controlled at the designed flow. This is achieved by

measuring the volumetric flow rate (FI 01) and transmitting it to the flow rate controller (FIC 01)

which controls the valve opening. This control loop operates as a reverse controller since the

valve opening should be decreased if the flow increases and is seen as a feedforward control

loop.

The ratios of the flow rates for the aqueous and organic phases through the stages are

controlled using a series of control loops as seen in Figure 7.8. All of these control loops

implement the ratio control method as described in Section 7.1.1. The organic to aqueous ratio

control for the scrubbing is described as an example. The set point for the volumetric flow rate

of the uranium loaded stream (organic) is transmitted to the flow rate controller (FIC 03) which

then controls the flow rate of the aqueous stream by using cascade control. These control loops

operate as direct controllers since the controlled flow rate should increase if the set point of the

wild flow rate increases and is seen as feedforward control loops.

The pH control in the respective stripping section stages is important and is difficult to control

due to the logarithmic dependence of the pH. The control for each of the stages is represented

by the single control loop in Figure 7.8. The pH is measured using an external pH probe

(PH 01), as discussed in Section 4.5, which does not come into contact with the organic phase

to prolong the lifetime of these probes. The measured pH is transmitted to a pH controller

(PHC 01) which then controls the flow of caustic to the respective stripping stage by

implementing cascade control. This control loop operates as a reverse controller since the

caustic flow rate should decrease if the pH increases and is seen as a feedback control loop.



7.1.4. Unit 4: Precipitation

The precipitation unit recovers the product from the uranium containing OK-liquor leaving the

solvent extraction unit. The product specifications for the ADU product are also achieved in this

unit to obtain optimum economical benefit. The volumetric flow rates in this unit are relatively

low since the concentration of uranyl sulphate complexes in the OK-liquor is high. To achieve

the desired product specifications and ensure adequate recovery, the relevant process

conditions are controlled within the design bounds. The following control objectives are

identified:

• Reactor temperature

• OK-liquor feed flow rate

• Reactor tank liquid levels

• Reactor liquid pH

• Ammonia gas make-up

Each of these process conditions are controlled to ensure the optimum operation of this unit.

The control schematic which shows the control loops for the two precipitation reactors is given

in Figure 7.9.



Figure 7.9: Control schematic of the precipitation reactors in the precipitation unit

In Figure 7.9 it is seen that the ADU is formed and precipitated in two precipitation reactors by

increasing the pH and ammonia gas addition. The OK-liquor is heated to the desired

temperature with steam in a heat exchanger before it is fed to the first precipitation reactor. The

uranyl sulphate complexes react with the ammonium fed to the reactor to form ADU which

precipitates at a higher pH. From here the solid ADU containing stream is sent to the solid-

liquid separation section.

The temperature of the precipitation reactors are controlled at a certain level to ensure the

reaction rate achieves the desired conversions of the uranyl sulphate complexes. The

temperature of the OK-liquor is controlled to manipulate the temperature in the reactors. To

achieve this the reactor temperature is measured (TI 01) and transmitted to the temperature

controller (TIC 01) which controls the flow of the steam to the heat exchanger by using a

cascade control loop. This control loop operates as a reverse controller since the steam flow

rate should be increased if the reactor temperature decreases. This control loop is seen as a

feedback control loop.



The OK-liquor feed flow rate together with the precipitation reactor liquid level determines the

residence time in the reactors. These parameters are therefore controlled to ensure sufficient

residence time which will achieve the desired conversions. The OK-liquor feed flow rate is

measured (FI 02) and transmitted to the flow rate controller (FIC 02) which controls the valve

opening. This control loop operates as a reverse controller since the valve opening should be

increased if the flow rate decreases. This control loop is seen as a feedforward control loop.

The reactor tank liquid levels are also controlled to ensure the desired residence time. This is

achieved by measuring the reactor liquid level (HI 01) in the first reactor and transmitting it to

the liquid level controller (HIC 01). The liquid level controller then controls the flow rate of the

exit stream from the first reactor by implementing a cascade control loop. The liquid level for the

second reactor is controlled similarly by measuring the liquid level (HI 02), transmitting the

height to the height controller (HIC 02) and controlling the flow rate of the exit stream from the

second reactor. These control loops operate as direct controllers since the exit flow rate should

be increased if the liquid level increases. These control loops are seen as feedback control

loops.

The pH levels of the reactors are controlled to achieve the desired reaction kinetics and ensure

precipitation of the ADU. The pH control in the first reactor is achieved by measuring the pH in

the reactor (PH 01) and transmitting it to the pH controller (PHC 01) which controls the ammonia

gas flow rate to the reactor. The ammonia flow rate is controlled by using a cascade control

loop. The pH control for the second reactor is done similarly. These control loops operate as

reverse controllers since the ammonia gas feed flow rate should be increased if the pH level

decreases. These control loops are seen as feedback control loops.

The ammonia gas is fed to the reactors as a mixture with air. To control the ammonia

concentration in the gas stream fed to the reactors, a ratio controller is implemented between

the air and ammonia gas feeds. This ratio control loop measures the volumetric flow rate of the

ammonia gas (FI 04) stream and transmits it to the ratio controller (FIC 01). The ratio controller

calculates a set point for the air volumetric flow rate by multiplying the flow rate of the ammonia

with the fixed ratio. The air flow rate is measured (FI 03) and transmitted to the ratio controller

(FIC 01) which controls the valve opening in the air feed stream according to the calculated set

point. This control loop operates as a direct controller since the valve opening in the air feed



stream should be increased if the flow rate of the ammonia gas feed stream increases. This

control loop is seen as a feedforward control loop.

The slurry stream containing the precipitated solid ADU is sent the solid-liquid separation

section where the final ADU product is produced to specification. The ADU solids are separated

from the liquids and washed to achieve the desired ADU purity and density. This section

consists of a thickener and three centrifuges. In this section it is important to control the flow

rates to ensure adequate separation and purification. The following control objectives are

identified for this section:

• Thickener liquid level

• ADU feed to centrifuges

• Wash water to the centrifuges

By controlling these process conditions the final ADU product is produced to specification. The

control schematic for the solid-liquid separation section in the precipitation unit is given in Figure

7.10.



Figure 7.10: Control schematic of the solid-liquid separation in the precipitation unit

In Figure 7.10 it is seen that the solids mass fraction of the ADU product is increased in a

thickener from where it is sent to two parallel centrifuges. In the centrifuges the ADU is washed

with water to remove remaining liquid impurities.



The liquid level in the thickener is controlled to ensure adequate residence time and effective

overflow. This is achieved by measuring the liquid level in the thickener (HI 01) and transmitting

it to the liquid level controller (HIC 01) which controls the flow rate of the thickener underflow

flow rate by using a cascade control loop. This control loop operates as a direct controller since

the thickener underflow rate should be increased if the liquid level in the thickener increases.

This control loop is seen as a feedback control loop.

The ADU streams to the two parallel centrifuges are controlled to ensure efficient operation of

the equipment. To achieve this, the flow rate of ADU is measured (FI 02 & FI 03) and

transmitted to the flow rate controllers (FIC 02 & FIC 04) which then control the valve position.

These control loops operate as reverse controllers since the valve opening should be increased

if the flow rate decreases. These control loops are seen as feedforward control loops.

The wash water flow rates used to wash the ADU product in the centrifuges are controlled at a

fixed ratio according to the ADU stream. Ratio control loops as discussed for the ammonia gas

feed make-up for the precipitation reactors are implemented to control the flow rate of the wash

water streams. These control loops operate as direct controllers since the wash water flow rate

should be increased if the ADU flow rate increases. These control loops are seen as

feedforward control loops.

The control strategies discussed above are used to control the many important process

conditions on the uranium extraction plant. These strategies are control loops and procedures

that each controls a certain process parameter. An integrated control system, that incorporate

and monitor all above control strategies, must be designed and implemented to achieve

effective and practical plant wide control.



7.2. Specific safety considerations for solvent extraction

To ensure the safety of the environment, equipment and employees, it is important to consider

several control aspects in the solvent extraction unit. These control aspects influence both the

economical and production performance. The influence of several process disturbances in the

solvent extraction unit are studied using a HAZOP level 3 analysis. The process disturbances

include no flow, low flow, high flow and several other disturbances. The HAZOP level 3 analysis

indicates shortcomings in the control strategies, which is rectified to reduce the safety risk of a

plant.

In this section a HAZOP level 3 analysis is done on the solvent extraction section. Before

commencing the HAZOP level 3 analysis, an fire hazard due to the extremely flammable

organic solvent is identified. The process control must also prevent solvent loss. The process

flow diagram used for which illustrates the different flow loops considered for the HAZOP level 3

analysis.



The five different liquid flow loops, seen in the process flow diagram, used for the HAZOP level

3 analysis are colour coded. The first liquid flow used is the aqueous phase flow in the

extraction section and is represented with the colour brown. The eluate stream is the uranium

containing feed stream to the solvent extraction section. The eluate flows into a tank which

supply capacitance to the system to reduce the effects of process disturbances. The aqueous

phase is contacted repeatedly with the organic solvent which extracts the uranium, before it is

recycled back to the leaching section.

The streams indicated as green, as seen in the process flow diagram, represent the organic

solvent flow loop. This loop flows counter-current in respect to the aqueous phase flows. The

organic flow enters the extraction section where the organic is loaded with uranium and

continues to the scrubbing section. At the scrubbing section impurities are removed from the

solvent before the uranium is stripped in the stripping section. Finally the organic solvent is

regenerated and recycled back to U03-ST03 (organic storage tank).

The demineralised water flow loop is indicated with the blue stream. The demineralised water is

stored in U03-ST04 from which it is continuously recycled through the scrubbing section to

remove impurities. When the demineralised water is saturated with impurities, it is drained from

the system to allow for fresh demineralised water.

The next flow loop is indicated with the orange stream flow and symbolises the aqueous phase

through the stripping section. This loop contains two feed streams i.e. the stripping solution and

caustic feed. The stripping solution flows counter-current in respect to the organic solvent and

removes the uranium from the organic solvent. The caustic is fed into each tank to control the

pH which has a significant effect on stripping efficiency.

The organic phase is regenerated by contacting it with caustic solution, which is represented

with the colour black. To achieve the desired concentrations the caustic and demineralised

water are mixed in a make-up tank (U03-MU02). The caustic solution is continuously recycled

through the regeneration section until the solution is saturated with impurities.

These five streams are used as an approach to the HAZOP level 3 analysis to simplify the

reasoning. The HAZOP level 3 analysis is show in Table 7.1 and Table 7.2.



Table 7.1: HAZOP 3 Proforma HS3A form

Path/line/node Identification

Eluate loop Organic loop Demin water

loop

OK liquor loop Regeneration

loop

Problem? No Yes Ref no.

No Yes Ref no.

No Yes Ref no.

No Yes Ref no.

No Yes Ref no.

High flow * 1 * 17 * 22 * 27 * 34

Low flow * 2 * 18 * 23 * 28 * 35

No flow * 3 * 19 * 24 * 29 * 36

Reverse flow * 4 * 4 * 4 * 4 * 4

High pressure * * * * *

Low pressure/vac * * * * *

High temperature * * * * *

Low temperature * * * * *

High level * 5 * 20 * 25 * 30 * 37

Low level * 6 * 21 * 26 * 31 * 38

High composition * * * * *

Low composition * * * * *

High pH * * * * 32 * 39

Low pH * * * * 33 * 40

Fast reaction, mix * * * * *

Slow reaction, mix * * * * *

High differential * 7 * 7 * 7 * 7 * 7

High stress * 8 * 8 * 8 * 8 * 8

Poor integrity * 9 * 9 * 9 * 9 * 9

Malfunction * 10 * 10 * 10 * 10 * 10

Impurities * 11 * 11 * 11 * 11 * 11

Lost containment * 12 * 12 * 12 * 12 * 12

Radiation * 13 * 13 * 13 * 13 * 13

Generation * 14 * 14 * 14 * 14 * 14

Maintenance * 15 * 15 * 15 * 15 * 15

Start/stop * * * * *

Emergency/ test * 16 * 16 * 16 * 16 * 16

Inoperability * * * * *



Table 7.2: HAZOP 3 record for Proforma HS3A

Ref. no:

Deviation Causes Consequences or Hazards Safeguards already provided

Recommendations or Actions

By

1 High flow − Upstream disturbances − Pump failure

− Overflow of tanks U03-ST01 and U03-ST02

− Flooding of settler − Loss of settling efficiency − Increase of organic flow

− Overflow weirs in settlers − Overflow valves on tanks

U03-ST01 and U03-ST02 − Back-up pumps

− High capacitance in tank U03-ST01

− Built in drain on U03-ST01 with level control

− Level control on tank U03-ST02

− Flow control on stream U03-AQ02

− Adequate alarm system

2 Low flow − Blockages in pipes − Pump failure

− Solvent loss − Loss of production − Phase inversion in settlers

− Capacitance due to tank U03-ST01

− Back-up pumps

− Filters − Emergency shut-down

procedure to ensure safety of solvent


3 No flow − Upstream shut-downs − Pump failure − Power failure − Blockages in pipes

− Loss of production − Pump cavitations − Up-stream down time


− Back-up pumps

− Emergency shut-down procedure to ensure safety of solvent

− High capacitance in tank U03-ST01


4 Reverse flow − Height difference − Pump damage − None − Non-return valves − Insure pumps are able to

handle reverse flow

5 High level − High flow up stream − Low flow downstream − No flow downstream

− Flooding of tanks and settlers

− Loss of product


U03-ST01 and U03-ST02

− Adequate alarm system − Level control on tank

U03-ST02

6 Low level − Low flow up stream − High flow downstream − No flow up stream

− Cavitation of pumps − Loss of production


− Emergency shut-down procedure




Ref. no:



By

7 High differential

− Valve position at start-up − Equipment damage − None − Adequate start-up procedure

− Pipe layout to reduce water hammer effect

8 High stress − Vacuum inside tanks − Internal collapse − Fire proof vent openings on tanks

− None

9 Poor integrity

− Corrosion − Spillage − Loss of raw materials − Down time

− Material of construction − Frequent inspections

10 Malfunction − Failure of instruments − Power failure − Instrument air − Pump failure − Mixer failure

− Loss of production − Unsafe working conditions − High, low or no flow

− Back-up pumps − Instrument redundancy − Alarm system on mixers

11 Impurities − Solids from upstream processes

− Impeller damage − Lower stage efficiency

− None − Filter

12 Lost containment

− Leaks − Evaporation

− Fire − Damage to environment − Unsafe working conditions

− Drainage systems − Enclosed tanks and

settlers

− Containment walls − Frequent inspections

13 Radiation − Nuclear radiation − Unsafe working conditions − Damage to environment

− Regular radiation testing personnel

− Correct PPE

− None

14 Generation − High flow rates − Half full pipes − Instrumentation short

circuits − Static − Personal accessories

(phones, etc.)

− Fire − Ensure flows lower than 1 m/s

− Air bleeds in pipe system − Fire proof instrumentation − Allow no personal

accessories on the SX plant

− Ground all equipment − Fibre grating cat walks



Ref. no:



By

15 Maintenance − Malfunction of instrumentation and equipment

− Corrosion − Removal of precipitated

solids from settlers

− Down time − Unsafe working conditions

− Frequent inspections − Operating procedure

16 Emergency/ test

− Fire − Equipment damage − Unsafe working conditions


− Frequent inspection − Adequate alarm system

17 High flow − Pump failure − Flooding of settlers − Solvent loss − Production loss

− Overflow weirs in settlers − Sump to recycle solvent

overflows − Back-up pumps − Overflow valve on tank

U03-ST03

− After settler to reduce solvent loss


− Flow control on streams U03-OG03 and U03-OG16

− Adequate alarm system − Built in drain on U03-

ST03 with level control

18 Low flow − Pump failure − Blockages in pipes

− Loss of production − Phase inversion

− Capacitance due to tan k U03-ST03

− Back-up pumps




19 No flow − Pump failure − Blockages in pipes − Power failure − Up- and downstream

shut-downs


− Back-up pumps − Capacitance due to tank

U03-ST03

− Filters − Adequate alarm system



Ref. no:



By

20 High level − High flow of recycle stream U04-OG16

− Low flow downstream − No flow downstream


− Loss of product


U03-ST03


U03-ST03

21 Low level − Low flow of recycle stream U04-OG16

− High flow downstream − No flow up stream

− Pump cavitations − Loss of production




22 High flow − Upstream disturbances at demin. water plant

− Pump failure

− Flooding of tank U03-ST04 and settlers

− Overflow weirs in settlers − Overflow valve on tank

U03-ST04 − Back-up pumps



23 Low flow − Upstream disturbances − Pump failure − Blockages in pipes

− Loss of production − Phase inversion


− Back-up pumps





shut-downs



U03-ST04


25 High level − High flow of upstream demin. water plant



− Loss of product


U03-ST04


U03-ST04 using the flow of U03-DW01



Ref. no:



By

26 Low level − Low flow up stream demin. water plant







− Flooding of tank U03-ST05, U03-ST06 and settlers

− Overflow weirs in settlers − Overflow valve on tanks

U03-ST05 and U03-ST06 − Back-up pumps


− Adequate alarm system − High capacitance in tank

U03-ST06 − Flow control on stream

U03-AQ09


− Loss of production − Phase inversion − Solvent loss


− Back-up pumps





shut-downs



U03-ST05


30 High level − High flow of upstream processes



− Loss of product


U03-ST05


U03-ST05 and U03-ST06

− Height control on tank U03-ST07 using stream U03-DW08



Ref. no:



By

31 Low level − Low flow up stream processes





− Adequate alarm system − Height control on tank

U03-ST07 using stream U03-DW08

32 High pH − pH of precipitation recycle − Caustic soda addition

− Precipitation of solids − Inadequate phase

separation

− 4 stripping settlers instead of 3

− Addition of caustic soda to stripping settler 2

− pH measurement at mixer instead of aqueous stream out

− pH control on tank U03-ST07

− Flow control on caustic soda addition to mixers

33 Low pH − pH of precipitation recycle − No flow of caustic soda

− Lower stage efficiency − 4 stripping settlers instead of 3

− Addition of caustic soda to stripping settler 2

− pH measurement at mixer instead of aqueous stream out

− Flow control on caustic soda addition to mixers

− pH control on tank U03-ST07


− Flooding of tank U03-MU02, U03-ST08 and settlers

− Overflow weirs in settlers − Overflow valve on tanks

U03-MU02 and U03-ST08 − Back-up pumps

− Level control on tank U03-MU02 using stream U03-DW07



− Loss of production − Phase inversion − Solvent loss

− Capacitance due to tank U03-MU02

− Back-up pumps






Ref. no:



By


shut-downs



U03-MU02


37 High level − High flow up stream processes


− Flooding of tanks and settler − Loss of product

− Overflow weirs in settler − Overflow valves on tanks

U03-MU02


U03-MU02 using stream U03-DW07

38 Low level − Low flow up stream processes



− Capacitance due to tank U03-MU02



U03-MU02 using stream U03-DW07

39 High pH − Caustic soda and Na2CO3 addition

− Inadequate phase separation

− None − pH measurement at mixer instead of aqueous stream out

− pH control on tank U03-MU02 using U03-ST09

40 Low pH − No flow of caustic soda and Na2CO3

− Lower stage efficiency − Inadequate solvent

regeneration

− None − pH measurement at mixer instead of aqueous stream out

− pH control in mixer using stream U03-S10



The following recommended actions are identified and rectified. From the HAZOP level 3

analysis the emergency shutdown and start-up procedures are revised to account for the

identified risks and to ensure adequate operating response. Filters are installed to remove alien

objects from the system which cause blockages. Non-return valves are used to protect

equipment against reverse flow. The pipe layout on the plant is specifically designed to reduce

the risk of damage if they experience a water hammer. Each tank on the solvent extraction site

is surrounded by a containment wall to prevent the fires from spreading and reduce spill

contaminated area. Due to the high acidic process conditions it is important to conduct frequent

inspections on equipment. It is necessary to remove all possible sources of ignition from the

solvent extraction plant and therefore fibre grated catwalks are used rather than metal catwalks

to prevent charge build-up.

After the HAZOP level 3 analysis is done, the recommended corrective actions are

implemented. In the following section the implementation procedure and discussion for these

corrective actions are given. Each of the protective measures is imperative for the continuous

efficient and safe operation of the plant.

7.3. Detailed process control for solvent extraction

The most important process parameters to control are the feed ratios of each section, feed flow

of the eluate to the extraction section and the pH in the stripping section. As identified in the

HAZOP level 3 analysis it is also important to control the level of the storage tanks to reduce the

probability of tank flooding. These levels are controlled by draining the tank to a sump when a

critical level is reached and therefore a draining system and sump is required. To prevent

unnecessary emergency shutdowns due to instrument failure, redundant instrumentation is

installed. A RAT (range, alarms and trip) list is composed for the implemented control loops to

ensure notification. All these identified control strategies are illustrated in the following process

flow diagram.



The process flow diagram shows twenty one control loops and extra specific indicators that

are installed in order to enable desired control of the solvent extraction section. According to

Svrcek et al. (2008) it is helpful to create a relative gain array to indicate the best control

scheme for the pairing of manipulated and controlled variables. The simplicity of the control

objectives and existing solvent extraction control strategies makes a relative gain array

unnecessary for this specific case. The control loops are discussed in five groups with

similar control objectives which are listed below:

• Basic flow control.

• Level control on tanks using the drain system.

• Level control on tanks using feed to the tank.

• pH control on make-up tanks.

• pH control on the mixer-settlers.

However, as mentioned certain instrumentation is installed to indicate malfunctioning of

equipment. For this reason electric current indicators are installed on all pumps and mixers,

and enables operators to switch any of this equipment on or off. These indicators allow

operators to observe the working status of these pumps and mixers from the control room.

Malfunctioning of the pump is shown by these indicators, for instance, if the pump is running

and a zero reading is shown, the pump is malfunctioning. Only selected electric current

indicators are shown in the process flow diagram to simplify the sketch.

7.3.1. Flow control loops

The basic flow control loops on selected streams are indicated with pink on the process flow

diagram. These flows are controlled at points where a certain flow rate is desired for

effective operation of the solvent extraction unit and other units. The flow control loops on

the streams entering the mixer-settlers are implemented to control the flow at a certain ratio

to the uranium containing phase flow through the mixer-settlers.



Basic flow control loops are fast responding with little dead time and no capacitance. It is

important to place the flow sensor upstream of the control valve to ensure more accurate

flow readings. Due to the no capacitance in this system, flow measurements are noisy and

therefore derivative action control is not implemented. PI controllers are widely used in the

industry for flow control and are simpler to tune. The quarter decay ratio is commonly not

observed for these control loops, due to the fast response of the flow controllers (Svrcek et

al., 2008: 147).

The simpler flow control loops control the flows of the raffinate recycle to the leaching and

the OK-liquor to the precipitation unit. These control loops control the flows to reduce

disturbances to the leaching and precipitation units. The tanks from which these streams

flow provide efficient capacitance to dampen disturbances.

The eluate feed to the extraction section is controlled using FIC 02 and is kept constant at a

set point. The organic phase flow fed to the extraction section is controlled with FIC 06 with

a set point calculated from the set point for FIC 02 multiplied with the specified ratio. The

demineralised water fed to the scrubbing section is similarly controlled with FIC 07 at a set

point which is calculated from FIC 06. The stripping solution flow entering the stripping

section is controlled with FIC 11 at a certain ratio to the set point of FIC 06. The caustic feed

to the regeneration section is controlled with FIC 20 which also has a set point at a certain

ratio to that of FIC 06. Finally the organic solvent recycle flow is controlled with FIC 21

which also has a set point at a certain ratio to that of FIC 06. Table 7.3 shows the

specifications for control loops discussed above.



Table 7.3: Specifications for flow controllers

Controller Description Process variable Set point

FIC 02 Controls eluate flow to

U03-MSE01.

Volumetric flow of

U03-AQ02. 25.816 m3/hr

FIC 04 Controls raffinate flow

to leaching.

Volumetric flow of

U03-AQ07.

1 x Set point of

FIC 02

FIC 06 Controls organic flow

to U03-MSE03.

Volumetric flow of

U03-OG03.

1.1 x Set point of

FIC 02

FIC 07

Controls demineralised

water flow to U03-

MSS03.

Volumetric flow of

U03-DW02.

0.2 x Set point of

FIC 06

FIC 11

Controls stripping

solution flow to U03-

MST01.

Volumetric flow of

U03-AQ09.

0.28 x Set point of

FIC 06

FIC 16 Controls OK-liquor flow

to precipitation.

Volumetric flow of

U03-AQ15.

1 x Set point of

FIC 11

FIC 20

Controls regeneration

solution flow to U03-

MSR01.

Volumetric flow of

U03-S10.

1 x Set point of

FIC 06

FIC 21

Controls organic

recycle flow to U03-

ST03.

Volumetric flow of

U03-OG16.

1 x Set point of

FIC 06

Controlling the flow and flow ratios are a simple and effective method to keep the mixer-

settler levels at the designed conditions. Therefore these flow control loops are important to

ensure effective operation of the mixer-settlers and prevent solvent loss.

7.3.2. Tank level control loops

There are two types of tank level control loops implemented in the solvent extraction unit.

One type is where the level is controlled by draining the tank, indicated with orange, while

the other controls the feed flow to the tank and is indicated with blue. The tanks supply

capacitance to the system to simplify control and dampen process disturbances. The

disturbances on the liquid surface may cause noise on the measured process variable and

therefore PID control is used with caution (Svrcek et al., 2008: 151).



The level on tanks U03-ST01, U03-ST02, U03-ST05 and U03-ST06 are controlled by

draining the tank to a sump. This control is implemented here, since it is important that the

tanks do not overflow. For this reason the liquid height itself is not controlled, but the tanks

is only drained at a dangerous liquid level and therefore on/off controllers are used in these

control loops. The draining system rather than the exit or feed streams are controlled to

reduce the influences of the level control on the greater system. The on/off control loops

make use of a dead band due to incapability of the controller to throttle the actuator. The

dead band used here has a maximum of 95 % and minimum of 90 % of the liquid level.

When the liquid level rises above the maximum point, the draining flow is opened and only

closed once the liquid level is below the minimum point of the dead band (Svrcek et al.,

2008: 94).

These controllers are cascade control loops, which is used seldom and malfunctions may

occur. Therefore it is important to have a flow indicator on the draining stream to ensure that

flow occurs when the valve is open. The specifications for these control loops are given in

Table 7.4.

Table 7.4: Specifications for drain system level control


HIC 01 Level control on

U03-ST01.

Liquid level of

U03-ST01.

On – 95 %

Off – 90 %


U03-ST02.

Liquid level of

U03-ST02.

On – 95 %

Off – 90 %


U03-ST05.

Liquid level of

U03-ST05.

On – 95 %

Off – 90 %


U03-ST06.

Liquid level of

U03-ST06.

On – 95 %

Off – 90 %

The level on tanks U03-ST04, U03-ST07 and U03-MU02 are controlled by the demineralised

water feed to the tanks. This is done to ensure adequate liquid levels and to simplify the pH

control on these tanks. The control loops implemented here are also cascade control loops

for more accurate control of the flow. The specifications for these control loops are shown in

Table 7.5.



Table 7.5: Specifications for feed flow level control



U03-ST04.

Liquid level of

U03-ST04. 75 %


U03-ST07.

Liquid level of

U03-ST07. 75 %


U03-MU02.

Liquid level of

U03-MU02. 75 %

The control loops of the liquid levels of these tanks create no disturbances that will influence

the solvent extraction process or downstream processes. Since the liquid levels of these

tanks are not of utmost importance to the system, their influences on the system are

minimised.

7.3.3. pH control loops

The pH is controlled in two equipment types in the solvent extraction unit, in the caustic

make-up tanks and in the mixers. The pH measurement and configuration of these two

control loops are different and are therefore discussed separately. To control the pH of any

system accurately, it is important to have a constant pH in the feed stream used to control it.

For this reason the pH on the make-up tanks are controlled to provide feed streams with a

stable pH.

The pH on tank U03-ST07 and U03-MU02 are controlled using similar control loops

represented by the green control loops. The caustic feed flows to these tanks are controlled

to control the pH in the tanks. The control loops implemented for this function is cascade

control loops to ensure more accurate control. It is important to note that the height of these

tanks are controlled using the demineralised water addition. These two control loops on

each tank are interdependent and therefore influence one another in the dynamic state.

However, this control strategy results in adequate control response to achieve the desired

control conditions. A simple case study shows that this control strategy is capable of

reaching the control objectives.



For the case study a situation where the liquid level is high and the pH is also high is

assumed and is a worst case scenario. In this case, the caustic feed to the tank is closed

and the demineralised water feed is also closed. The liquid level will soon drop due to a

constant exit stream, causing the demineralised water feed to open slowly, which will

decrease the pH. Quicker response time for any other situation is expected.

The pH of the mixers is simply controlled using a cascade control loops on the caustic flows

indicated by the yellow control loops. The measurement of the pH in the mixers is

complicated since the mixers contain a dispersion of aqueous and organic phases. Direct

pH measurement of the dispersion is possible, from which the aqueous pH is derived.

However, the pH probes used in the direct measurement comes in contact with the organic

phase which significantly reduces the lifetime of the costly pH probes. For this reason, a

small settling box is used (as described in Section 4.5), where settling takes place, and the

pH of the aqueous phase is measured. This vessel must have a small volume to reduce the

residence time and ultimately reduce the dead time on the control loop. The specifications

for these control loops are given in Table 7.6.

Table 7.6: Specifications for pH control


PHC 10 pH control on

U03-ST07. pH of U03-ST07. 9.5


U03-MST02. pH of U03-MST02. 5.1


U03-MST03. pH of U03-MST03. 4.8


U03-MST04. pH of U03-MST04. 3.9


U03-MU02. pH of U03-MU02. 9.5

The pH control on the mixers in the stripping section is important to ensure adequate stage

efficiency and phase separation. The pH in the in the first mixer of the stripping section is

high due to the stripping solution recycle from the precipitation unit. Thereafter, the pH is

controlled in a decreasing order in the second, third and fourth mixers, as indicated in Table

7.6.



7.3.4. Range, alarms and trips

The HAZOP level 3 analysis done on the solvent extraction section shows that alarm

systems and trips are required to ensure plant safety. For this reason a RAT list is compiled

and shown in Table 7.7.

Table 7.7: RAT list for the solvent extraction section

Loop number

Tag number

Instrument type

Service RAT information

Range Alarm Trip

HIC 01

HI 01 Ultra sonic

level

Indicate level on U03-

ST01 0-100%

H: 95%

L: 20%

HH:99%

LL: 5%

FI 01 Flow meter Indicates flow to U03-

MSE01

0-35

m3/hr

FIC 02 FI 02 Flow meter Indicates flow to U03-

MSE01

0-35

m3/hr

H: 26.5

L:24.5

HH:28.5

LL: 22

AI 03-

25 Amp meter

Indicates amp of the

respective pump or

mixer

0-20A H:12

L: 3

HH: 15

LL: 0

FIC 04 FI 04 Flow meter Indicates flow to

leaching

0-35

m3/hr

H: 26.5

L:24.5

HH:28.5

LL: 22

HIC 05

HI 05 Ultra sonic

level


ST01 0-100%

H: 95%

L: 20%

HH:99%

LL: 5%

FI 05 Flow meter Indicates flow to

Drain

0-35

m3/hr


MSE03

0-35

m3/hr

H: 29.5

L:27.5

HH:31

LL: 25


MSS03

0-35

m3/hr

H: 8.5

L:7

HH:9.5

LL: 6



HIC 08

HI 08 Ultra sonic

level


ST04 0-100%

H: 80%

L: 20%

HH:95%

LL: 5%


ST04

0-35

m3/hr

HIC 09

HI 09 Ultra sonic

level


ST07 0-100%

H: 80%

L: 20%

HH:95%

LL: 5%


ST07

0-35

m3/hr

PHC 10

PH 10 pH meter Indicate pH on U03-

ST07 1-14pH

H: 9

L: 6

HH:10.5

LL: 4.5


ST07

0-35

m3/hr


MST01

0-35

m3/hr

H:8.5

L:7

HH: 9.5

LL: 6

HIC 12

HI 12 Ultra sonic

level


ST07 0-100%

H: 95%

L: 20%

HH:99%

LL: 5%


ST07

0-35

m3/hr

PHC

13-15

PH pH meter Indicates pH of U03-

ST07 1-14pH

H: 8

L: 3

HH:9.5

LL: 2

FI Flow meter Indicates flow to U03-

ST07

0-35

m3/hr


ST07

0-35

m3/hr

H:8.5

L:7

HH:9.5

LL:6

HIC 17

HI 17 Ultra sonic

level


ST06 0-100%

H: 95%

L: 20%

HH:99%

LL: 5%


ST06

0-35

m3/hr

PHC 18

PH 18 pH meter Indicates pH of U03-

MU02 1-14pH

H: 8

L: 3

HH:9.5

LL: 2


MU02

0-35

m3/hr

HIC 19

HI 17 Ultra sonic

level


MU02 0-100%

H: 80%

L: 20%

HH:95%

LL: 5%

FI 17 Flow meter Indicates flow from

U03-MU02

0-35

m3/hr



PHC 20

PH pH meter Indicates pH of U03-

MSR01 1-14pH

H: 8

L: 3

HH:9.5

LL: 2

FI Flow meter Indicates flow to U03-

MSR01

0-35

m3/hr


MSE03

0-35

m3/hr

H: 29.5

L:27.5

HH:31

LL: 25

Installing these alarms and trips will safeguard the equipment and effectively notify plant

personnel of the process disturbances. The overall plant safety and efficiency is greatly

enhanced by the implementation of these guidelines.

7.4. Dynamic control analysis

In this section the dynamic control response is simulated and analysed. This gives a greater

understanding of the dynamic behaviour of the implemented control system. The solvent

extraction unit is simulated using Aspen HYSYS®. It is decided to simulate the different

sections separately to simplify the flowsheet which contains recycle streams. The flow and

pH control loops are analysed and optimised in this section.

HYSYS does not simulate open tanks which produced several obstacles in the simulation of

the effectively open tank system. For this reason air is included in the component list and

streams to simulate the open tanks. Each mixer-settler unit is represented with a three

phase separator with the same volume and footprint area as that designed for the settlers.

In the design, the organic phase is 93 % kerosene, and therefore the organic phase is

represented by 100 % C12H24 in the simulation. The aqueous phase is always represented

by water.

For the desired separation to take place the thermodynamic model data in the simulation is

required. The NRTL thermodynamic package is used to simulate the system. The binary

activity coefficients for the liquid-liquid equilibrium are acquired using Aspen data bank. The

binary coefficients between the air and both liquids are set to zero to ensure no

thermodynamic interactions.



The HYSYS simulation is specifically used to simulate the liquid flow and the storage tank

liquid levels. Since this is a complex system and HYSYS does not have specialised process

blocks to simulate all the different equipment, only the above mentioned process parameters

are simulated.

7.4.1. Tuning control loops

To achieve effective control with the control loops, it is tuned to give satisfactory response

time with limited overshoot. By tuning the control, the control efficiency is optimised and the

control equipment, which includes control valves, is protected. An example of this is the

controllers in the flow control loops, which are PI-controllers instead of PID-controllers to

reduce the noise on control valves.

There are three parameters in a PID-controller that determines the control aggressiveness of

the controller. The gain (Kc), the integral time (Ti) and the differential time (Td) are tuned or

varied to obtain the desired control action on the specific system. HYSYS does include an

auto-tune function which automatically derive specific values for the controller parameters

(Kc, Ti and Td) to achieve satisfactory control (Svrcek et al., 2008: 107-110).

There are also methods, such as the Ziegler-Nichols open-loop method, available to obtain

these parameters from the simulated process (Svrcek et al., 2008: 125). Because of the

complex system and all the control equipment such as pumps, steady state is not easily

reached in the solvent extraction simulation. Therefore the Ziegler-Nichols method is not

implemented on the simulation, since it requires the simulation to reach steady state. In this

section the auto-tune function in HYSYS is used to optimise the control action of each

control loop.

In Section 7.3.1 it is mentioned that several of the flow control loops are ratio control loops.

The tuning of these ratio control loops are described by taking the flow ratio control in the

extraction section as an example. The simulation flowsheet in HYSYS for the extraction

section is shown in Figure 7.11.



Figure 7.11: Extraction simulation flowsheet

In Figure 7.11, the flow controllers that are using ratio control are indicated with the red

circles, where the eluate flow on the left is the independent stream and the organic flow on

the right is the dependent stream. The set point of the organic feed controller (FIC 06) is

directly calculated from the set point of the eluate flow controller (FIC 02). This is different

from the ratio control methods discussed by Svrcek et al. (2008) in that set point of the

independent stream, rather than the process variable, is measured to calculate the set point

of the dependent stream. This is done to ensure that process disturbances in the eluate flow

do not cause disturbances in the organic feed to the extraction section. Since the flow

control action of the flow control loops are fast and effective, the desired ratio is achieved

quickly. Figure 7.12 shows the eluate and organic flow into the extraction section when a

step change in the set point of the eluate feed is made.



Figure 7.12: Ratio control results in extraction section

The intended ratio is 1.1:1 for the organic to eluate flow rate into the system. From Figure

7.14 it is calculated that the ratio remains 1.1:1 before and after the change in the set point

of the eluate feed. Figure 7.14 shows that the implemented ratio control system is effective.

As an example of the tuning procedure followed to auto-tune the liquid level control loops in

the simulation, the liquid level controller (HIC 08) on U03-ST04 is tuned to optimise the

control action. This is a cascade control loop with the liquid level controller as the primary

controller and the flow controller as the secondary controller. The simulation flowsheet for

the scrubbing section in HYSYS is shown in Figure 7.13.



Figure 7.13: Scrubbing simulation flowsheet

Figure 7.13 shows the cascade control loop in the red circle which is used to control the

liquid level in the tank. In a cascade control loop it is important to ensure that the secondary

control loop has a faster (approximately four times faster) response time than that of the

primary control loop. If this is achieved the primary controller is not influenced by the control

of the secondary controller and therefore cascade control loop will function properly. Since

this control loop is a cascade control loop it is tuned by following the procedure below

(Svrcek et al., 2008: 134).

• Place the primary control loop in the manual control mode.

• Tune the secondary control loop as an independent loop.

• Place the primary control loop back in the automatic control mode.

• Tune the primary control loop normally.

Using this procedure, the cascade control loop is tuned and the effect is simulated. Figure

7.14 shows the liquid level control loop results.



Figure 7.14: The primary control loop results

Figure 7.14 shows the liquid level in the tank as the black line, while the set point for the

liquid level is indicated by the red line. When the set point is increased with 5 %, the control

response is fast, with an overshoot of approximately 39 % and a decay ratio of 0.08. The

high overshoot is not desired, but the small decay ratio shows adequate control. The rise

time is approximately 4 minutes and the response time 64 minutes. When the set point of

the liquid level is decreased with 5 %, a slower control action is observed. The overshoot for

this step is 27 % with a decay ratio of 0.18. The rise time for the downward step is 28

minutes with a response time of 64 minutes. The less aggressive control action of the

downward step is observed due to the limit on the exit stream flow caused by the pump. In



Figure 7.15 the set point and process variable (feed flow rate) of the secondary control loop

is shown for the same action as in Figure 7.14.

Figure 7.15: The secondary control loop results.

Figure 7.15 shows that the process variable in the secondary control loop follows the varying

set point closely and this implies a fast control action. This results in the effective control

action of the primary loop seen in Figure 7.14. All the control loops seen in Figures 7.11 and

7.15 are tuned using the auto-tuner function in HYSYS to simplify and aid the simulation. A

list of the controller parameters obtained for each of these control loops are given in Table

7.8.



Table 7.8: Control loop tuning values

Control loop Description KC Ti TD

Extraction

FIC-100 Level control on tank U03-ST02 0.10000 0.01000 -

FIC-101 OK liquor feed control to U03-MSE01 0.04860 0.00896 -

FIC-102 Organic feed control to U03-MSE03 0.05890 0.01120 -

LIC-100 Level control on tank U03-ST01 62.10000 1.31000 -

LIC-102 Level control on tank U03-ST03 1.00000 0.01000 -

Scrubbing

FIC-100 Control demin water feed to U03-MSS03 0.14400 0.00852 -

FIC-101 Control loaded organic feed to U01-MSS01 0.03940 0.00900 -

FIC-102 Cascade level control on tank U03-ST04 0.11800 0.01090 -

LIC-102 Cascade level control on tank U03-ST04 14.80000 1.47000 0.32600

After all these control loops are tuned, the simulation will run smoothly and possible

modifications are analysed. With all the short-comings of the HYSYS simulation it still

serves as a good indication of the validity of the control strategies.

7.4.2. Variable pairing

In the control of the solvent extraction unit there are several systems with multi-variable

inputs to control the process conditions. These single-loop control schemes may interact

and therefore influence each other. The combination of variables that results in the best

overall control is the combination of control loops that have the smallest influence on each

other. In many cases this pairing is obtained through logic reasoning, however there are

methods to indicate the best combination of variable pairing (Svrcek et al., 2008: 215).

The control system on the caustic storage tank (U03-ST07), used to control the pH in the

mixers, has a liquid level and pH control loop on the tank. It is decided to control these

parameters using the caustic and demineralised water feed streams to the tank. As an

example of variable pairing, these control loops are analysed to obtain the optimum variable

pairing. Before the variable pairing is done, the system is simulated with an appropriate

mathematical model in Excel. The HYSYS simulations are not used since the available

software package is incapable of simulating electrolyte systems.



In this simulation the height of the liquid level and the pH of the tank contents are calculated

as functions of the caustic and demineralised water feed. In this simulation the following

assumptions are made:

• Steady state conditions

• 1.5 mole/L NaOH in caustic feed

• Kw = 10-14

• Demineralised water pH = 7

The system is then solved to satisfy mass balances and equilibrium conditions. The variable

pairing is obtained by determining the relative gain array (RGA) for the system. This is done

by increasing the demineralised water volumetric feed to the tank with 20 % and recording

the pH and liquid level change. This is also done for a 20 % change in the caustic feed.

From these results a RGA is determined and given in Table 7.9 (Svrcek et al., 2008: 218-

219).

Table 7.9: Relative gain array for control on U03-ST07

Caustic feed Demineralised water feed

pH 1 0 Liquid level 0 1

From Table 7.9 it is clear that the pH of the tank must be controlled with the caustic feed and

the liquid level in the tank must be controlled with the demineralised water feed. In this

control scheme the two control loops has a minimal influence on each other. This is

expected from logical reasoning, since the demineralised water feed stream to the tank has

a little influence on the pH compared to the caustic feed. The Niederlinski index described

by Svrcek et al., (2008) estimates the stability of a control scheme in the dynamic state. The

Niederlinski index for this system is calculated as one (positive), which concludes that this

variable pairing control scheme is stable in dynamic state.



7.5. Conclusion

The control strategies for the plant wide control for the uranium extraction plant are able to

control all important process parameters. These strategies maintain the desired process

conditions and ensure the safety of the plant personnel, equipment and the environment.

The detail discussion of the control strategies for the solvent extraction unit shows that the

intended control is sufficient and also enables the optimisation of the process.

The simulation of the control system on the solvent extraction unit in HYSYS gave a better

insight into specific obstacles in the control. From this simulation it is for instance decided to

increase the volume of specific storage tanks in the solvent extraction unit to provide greater

capacitance to the system. This control system for the uranium extraction plant is a

conceptual system which serves as a starting point for further work to create a proper

functioning control system.


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229 Appendix A: Mass and energy balances

Appendix A: Mass and energy balances

A.1. Mass balance

The individual mass balances for the separate units are given in Appendix A. In these mass balances all the materials present in the

entire system is tabulated and the mass flows of the materials that is present in the specific unit is given in Table A.2 to A.7. The

stream names given correspond to the stream names used in the Aspen Tech® simulation and a short description of each stream is

given after the table. This simulation is shown in Figure A.1.

A summary of all the process units is given in Table A.1 with the mass flow in and out of the unit. As seen in Table A.1 the errors are

very small which is an indication that the inflow and outflow of the processes match and the law of conservation is confirmed.

Table A.1: Summary of mass balances for individual units

Process unit Mass flow in (kg/hr) Mass flow out (kg/hr) % error

Leaching 925 163 925 163.1 3.68E-08

CCD 1 526 960 1 526 960 6.044E-08

Neutralization 992 488 992 487.8 1.01E-08

Ion exchange 629 626 629 612.1 2.14E-03

Solvent extraction 41 815.8 41 836.78 5.01E-02

Precipitation 10 355.1 10 355.08 8.94E-06



Figure A.1: Aspen Tech® simulation



Table A.2: Mass balance for leaching process

IN

OUT Mass flow kg/hr

Mass flow kg/hr

ORE HNO3 H2SO4 LOAD

LEACHPRO

880 000 2 498.15 15 323.52 27 341.5

925 163.1

Liquids Solids Liquids Solids Liquids Solids Liquids Solids

Liquids Solids

H2O 350000 0 648.148 0 0.055 0 24064 0

380110 0

SIO2 0 379149 0 0 0 0 0 0

0 383824

MUSCV 0 54550 0 0 0 0 0 0

0 51167.9

CLINOCL 0 10802 0 0 0 0 0 0

0 4320.79

PYROPHLT 0 52389.6 0 0 0 0 0 0

0 49770.1

PYRITE 0 7021.28 0 0 0 0 0 0

0 6810.65

ALBITE 0 25924.7 0 0 0 0 0 0

0 25665.5

UO2 0 75.614 0 0 0 0 0 0

0 0.03025

UTI2O6 0 70.213 0 0 0 0 0 0

0 30.8936

UCLPO6 0 5.401 0 0 0 0 0 0

0 0.10802

UO2SIO2 0 10.802 0 0 0 0 0 0

0 0.21604

UO2-2W 0 1.08 0 0 0 0 0 0

0 0.0216

U3O8 0 0 0 0 0 0 0 0

0 0

TIO2 0 0 0 0 0 0 0 0

0 14.6155

K+ 0 0 0 0 0 0 0 0

331.985 0

AL+++ 0 0 0 0 0 0 0 0

2233.55 0

MG++ 0 0 0 0 0 0 0 0

507.751 0



FE++ 0 0 0 0 0 0 0 0

1166.72 0

FE+++ 0 0 0 0 0 0 0 0

0 0

FESO4- 0 0 0 0 0 0 74.527 0

509.879 0

NA+ 0 0 0 0 0 0 0 0

22.7283 0

UO2++ 0 0 0 0 0 0 0 0

10.3293 0

US2O10-- 0 0 0 0 0 0 0 0

176.491 0

US3O14-- 0 0 0 0 0 0 4.785 0

4.78546 0

H3O+ 0.001 0 159.719 0 323.55 0 982.086 0

165.462 0

OH- 0.001 0 0 0 0 0 0 0

1.84E-08 0

CL- 0 0 0 0 0 0 0 0

0.4686 0

PO4--- 0 0 0 0 0 0 0 0

1.25531 0

H2SO4 0 0 0 0 13348.8 0 0 0

2.30E-08 0

HSO4- 0 0 0 0 1651.07 0 0 0

2526.22 0

SO4-- 0 0 0 0 0 0 2216.1 0

14351.3 0

HNO3 0 0 1169.67 0 0 0 0 0

0.03813 0

HNO2 0 0 0 0 0 0 0 0

0 0

NO3- 0 0 520.612 0 0 0 0 0

1076.51 0

NO2 0 0 0 0 0 0 0 0

37.1426 0

NO 0 0 0 0 0 0 0 0

263.728 0

R-NO3 0 0 0 0 0 0 0 0

0 0

R2-UC 0 0 0 0 0 0 0 0

0 0

R2-SO4 0 0 0 0 0 0 0 0

0 0

R-FEC 0 0 0 0 0 0 0 0

0 0



L4US3O14 0 0 0 0 0 0 0 0

0 0

L2-SO4 0 0 0 0 0 0 0 0

0 0

KEROSENE 0 0 0 0 0 0 0 0

0 0

ISODEC 0 0 0 0 0 0 0 0

0 0

NH3 0 0 0 0 0 0 0 0

0 0

AMMON-01 0 0 0 0 0 0 0 0

0 0

NH4+ 0 0 0 0 0 0 0 0

0 0

ADU 0 0 0 0 0 0 0 0

0 0

H4SIO 0 0 0 0 0 0 0 0

61.4052 0

NH4NO3 0 0 0 0 0 0 0 0

0 0

LIME 0 0 0 0 0 0 0 0

0 0

CA++ 0 0 0 0 0 0 0 0

0 0

CAOH+ 0 0 0 0 0 0 0 0

0 0

CA(OH)2 0 0 0 0 0 0 0 0

0 0

CALCI(S) 0 0 0 0 0 0 0 0

0 0

AL(OH)3 0 0 0 0 0 0 0 0

0 0

FE(OH)2 0 0 0 0 0 0 0 0

0 0

MG(OH)2 0 0 0 0 0 0 0 0

0 0

Total 350000 530000 2498.15 0 15323.5 0 27341.5 0

403558 521605

Ore = The ore feed received from the mine, already mixed with water Leachpro = Leach product

HNO3 = 68% Nitric acid feed

H2SO4 = 98% Sulphuric acid feed

Load = Solvent extraction recycle



Table A.3: Mass balance for CCD process

IN

OUT

Mass flow kg/hr

Mass flow kg/hr

LEACHPRO 6 21

SOLIDS PREG-OF

925 163.14 601 796 0.338

924 903.6 602 056.5

Liquids Solids Liquids Solids Liquids Solids

Liquids Solids Liquids Solids

H2O 380110.45 0 567211.1 0 0.998 0

380121.2 0 567186.9 0

SIO2 0 383824.16 0 0.002 0 0

0 383824.1 0 0.045

MUSCV 0 51167.877 0 0 0 0

0 51167.87 0 0.006

CLINOCL 0 4320.79 0 0 0 0

0 4320.79 0 0.001

PYROPHLT 0 49770.101 0 0 0 0

0 49770.1 0 0.006

PYRITE 0 6810.645 0 0 0 0

0 6810.645 0 0.001

ALBITE 0 25665.493 0 0 0 0

0 25665.49 0 0.003

UO2 0 0.03 0 0 0 0

0 0.03 0 0

UTI2O6 0 30.894 0 0 0 0

0 30.894 0 0

UCLPO6 0 0.108 0 0 0 0

0 0.108 0 0

UO2SIO2 0 0.216 0 0 0 0

0 0.216 0 0

UO2-2W 0 0.022 0 0 0 0

0 0.022 0 0

U3O8 0 0 0 0 0 0

0 0 0 0

TIO2 0 14.616 0 0 0 0

0 14.616 0 0

K+ 331.985 0 495.384 0 0 0

331.985 0 495.384 0

AL+++ 2233.547 0 3332.871 0 0 0

2233.547 0 3332.871 0

MG++ 507.751 0 757.66 0 0 0

507.751 0 757.66 0



FE++ 1166.719 0 1740.963 0 0 0

1166.719 0 1740.963 0

FE+++ 0 0 0 0 0 0

0 0 0 0

FESO4- 509.879 0 637.268 0 0 0

427.07 0 720.077 0

NA+ 22.728 0 33.915 0 0 0

22.728 0 33.915 0

UO2++ 10.329 0 0.001 0 0 0

0.001 0 9.828 0

US2O10-- 176.491 0 0.018 0 0 0

0.012 0 177.357 0

US3O14-- 4.785 0 0 0 0 0

0 0 4.785 0

H3O+ 165.462 0 269.336 0 0 0

180.496 0 269.605 0

OH- 0 0 0 0 0 0

0 0 0 0

CL- 0.469 0 0.699 0 0 0

0.469 0 0.699 0

PO4--- 1.255 0 1.873 0 0 0

1.255 0 1.873 0

H2SO4 0 0 0 0 0 0

0 0 0 0

HSO4- 2526.22 0 3655.107 0 0 0

2449.504 0 3653.731 0

SO4-- 14351.285 0 21603.75 0 0 0

14477.9 0 21554.06 0

HNO3 0.038 0 0.055 0 0 0

0.037 0 0.055 0

HNO2 0 0 0 0 0 0

0 0 0 0

NO3- 1076.514 0 1606.363 0 0 0

1076.516 0 1606.363 0

NO2 37.143 0 55.424 0 0 0

37.143 0 55.424 0

NO 263.728 0 393.531 0 0 0

263.728 0 393.531 0

R-NO3 0 0 0 0.004 0 0

0 0.004 0 0

R2-UC 0 0 0 0.062 0 0

0 0.062 0 0

R2-SO4 0 0 0 0.547 0 0

0 0.546 0 0

R-FEC 0 0 0 0.024 0 0

0 0.024 0 0



L4US3O14 0 0 0 0 0 0

0 0 0 0

L2-SO4 0 0 0 0 0 0

0 0 0 0

KEROSENE 0 0 0 0 0 0

0 0 0 0

ISODEC 0 0 0 0 0 0

0 0 0 0

NH3 0 0 0 0 0 0

0 0 0 0

AMMON-01 0 0 0 0 0 0

0 0 0 0

NH4+ 0 0 0 0 0 0

0 0 0 0

ADU 0 0 0 0 0 0

0 0 0 0

H4SIO 61.405 0 0 0 0 0

0 0 61.405 0

NH4NO3 0 0 0 0 0 0

0 0 0 0

LIME 0 0 0 0 0 0

0 0 0 0

CA++ 0 0 0 0 0 0

0 0 0 0

CAOH+ 0 0 0 0 0 0

0 0 0 0

CA(OH)2 0 0 0 0 0 0

0 0 0 0

CALCI(S) 0 0 0 0 0 0

0 0 0 0

AL(OH)3 0 0 0 0 0 0

0 0 0 0

FE(OH)2 0 0 0 0 0 0

0 0 0 0

MG(OH)2 0 0 0 0 0 0

0 0 0 0

Total 403558.19 521604.95 601795.4 0.639 0.998 0

403298.1 521605.5 602056.4 0.061

Leachpro = Leach product Solids = Solids to neutralization

6 = Barren liquor Preg-of = Pregnant leach liquor

21 = Wash water make-up



Table A.4: Mass balance for neutralization process

IN

OUT

Mass flow kg/hr

Mass flow kg/hr

SOLIDS LIME

GOLD

924 903.6 67 584.13

992 487.8

Liquids Solids Liquids Solids

Liquids Solids

H2O 380121.2 0 57398 0

435043.3 0

SIO2 0 383824.1 0 0

0 383824.1

MUSCV 0 51167.87 0 0

0 51167.87

CLINOCL 0 4320.79 0 0

0 4320.79

PYROPHLT 0 49770.1 0 0

0 49770.1

PYRITE 0 6810.645 0 0

0 6810.645

ALBITE 0 25665.49 0 0

0 25665.49

UO2 0 0.03 0 0

0 0.03

UTI2O6 0 30.894 0 0

0 30.894

UCLPO6 0 0.108 0 0

0 0.108

UO2SIO2 0 0.216 0 0

0 0.216

UO2-2W 0 0.022 0 0

0 0.022

U3O8 0 0 0 0

0 0

TIO2 0 14.616 0 0

0 14.616

K+ 331.985 0 0 0

331.985 0

AL+++ 2233.547 0 0 0

0 0

MG++ 507.751 0 0 0

46.205 0



FE++ 1166.719 0 0 0

0 0

FE+++ 0 0 0 0

0 0

FESO4- 427.07 0 0 0

427.07 0

NA+ 22.728 0 0 0

22.728 0

UO2++ 0.001 0 0 0

0.001 0

US2O10-- 0.012 0 0 0

0.012 0

US3O14-- 0 0 0 0

0 0

H3O+ 180.496 0 0 0

0 0

OH- 0 0 0 0

4.411 0

CL- 0.469 0 0 0

0.469 0

PO4--- 1.255 0 0 0

1.255 0

H2SO4 0 0 0 0

0 0

HSO4- 2449.504 0 0 0

0 0

SO4-- 14477.9 0 0 0

16901.98 0

HNO3 0.037 0 0 0

0 0

HNO2 0 0 0 0

0 0

NO3- 1076.516 0 0 0

1076.552 0

NO2 37.143 0 0 0

37.143 0

NO 263.728 0 0 0

263.728 0

R-NO3 0 0.004 0 0

0 0.004

R2-UC 0 0.062 0 0

0 0.062

R2-SO4 0 0.546 0 0

0 0.546

R-FEC 0 0.024 0 0

0 0.024



L4US3O14 0 0 0 0

0 0

L2-SO4 0 0 0 0

0 0

KEROSENE 0 0 0 0

0 0

ISODEC 0 0 0 0

0 0

NH3 0 0 0 0

0 0

AMMON-01 0 0 0 0

0 0

NH4+ 0 0 0 0

0 0

ADU 0 0 0 0

0 0

H4SIO 0 0 0 0

0 0

NH4NO3 0 0 0 0

0 0

LIME 0 0 0 10186.13

0 0

CA++ 0 0 0 0

7272.343 0

CAOH+ 0 0 0 0

10.529 0

CA(OH)2 0 0 0 0

0 0

CALCI(S) 0 0 0 0

0 0

AL(OH)3 0 0 0 0

0 6457.574

FE(OH)2 0 0 0 0

0 1877.368

MG(OH)2 0 0 0 0

0 1107.526

Total 403298.1 521605.5 57398 10186.13

461439.8 531048

Solids = Solids to neutralization Gold = Solids to gold plant

Lime = Lime used for neutralization



Table A.5: Mass balance for ion exchange process

IN

OUT

Mass flow kg/hr

Mass flow kg/hr

PREG-OF RES-MU ELUANT WATERSIO 6 ELUATE SIO2-OUT

602 056.5 15.12 27 528.58 25.35 601 796 27 752.28 63.78

Liquids Solids Liquids Solids Liquids Liquids

Liquids Solids Liquids Liquids

H2O 567186.9 0 1 0 24451.18 25.352

567211.1 0 24451.99 2.307

SIO2 0 0.045 0 0 0 0

0 0.002 0 0

MUSCV 0 0.006 0 0 0 0

0 0 0 0

CLINOCL 0 0.001 0 0 0 0

0 0 0 0

PYROPHLT 0 0.006 0 0 0 0

0 0 0 0

PYRITE 0 0.001 0 0 0 0

0 0 0 0

ALBITE 0 0.003 0 0 0 0

0 0 0 0

UO2 0 0 0 0 0 0

0 0 0 0

UTI2O6 0 0 0 0 0 0

0 0 0 0

UCLPO6 0 0 0 0 0 0

0 0 0 0

UO2SIO2 0 0 0 0 0 0

0 0 0 0

UO2-2W 0 0 0 0 0 0

0 0 0 0

U3O8 0 0 0 0 0 0

0 0 0 0

TIO2 0 0 0 0 0 0

0 0 0 0

K+ 495.384 0 0 0 0 0

495.384 0 0 0

AL+++ 3332.871 0 0 0 0 0

3332.871 0 0 0

MG++ 757.66 0 0 0 0 0

757.66 0 0 0



FE++ 1740.963 0 0 0 0 0

1740.963 0 0 0

FE+++ 0 0 0 0 0 0

0 0 0 0

FESO4- 720.077 0 0 0 0 0

637.268 0 82.808 0

NA+ 33.915 0 0 0 0 0

33.915 0 0 0

UO2++ 9.828 0 0 0 0 0

0.001 0 75.942 0

US2O10-- 177.357 0 0 0 0 0

0.018 0 68.141 0

US3O14-- 4.785 0 0 0 0 0

0 0 0 0

H3O+ 269.605 0 0 0 575.859 0

269.336 0 575.006 0

OH- 0 0 0 0 0 0

0 0 0 0

CL- 0.699 0 0 0 0 0

0.699 0 0 0

PO4--- 1.873 0 0 0 0 0

1.873 0 0 0

H2SO4 0 0 0 0 0 0

0 0 0 0

HSO4- 3653.731 0 0 0 2073.462 0

3655.107 0 2077.817 0

SO4-- 21554.06 0 0 0 428.076 0

21603.75 0 420.579 0

HNO3 0.055 0 0 0 0 0

0.055 0 0 0

HNO2 0 0 0 0 0 0

0 0 0 0

NO3- 1606.363 0 0 0 0 0

1606.363 0 0 0

NO2 55.424 0 0 0 0 0

55.424 0 0 0

NO 393.531 0 0 0 0 0

393.531 0 0 0

R-NO3 0 0 0 0 0 0

0 0.004 0 0

R2-UC 0 0 0 0 0 0

0 0.062 0 0

R2-SO4 0 0 0 14.123 0 0

0 0.547 0 0

R-FEC 0 0 0 0 0 0

0 0.024 0 0



L4US3O14 0 0 0 0 0 0

0 0 0 0

L2-SO4 0 0 0 0 0 0

0 0 0 0

KEROSENE 0 0 0 0 0 0

0 0 0 0

ISODEC 0 0 0 0 0 0

0 0 0 0

NH3 0 0 0 0 0 0

0 0 0 0

AMMON-01 0 0 0 0 0 0

0 0 0 0

NH4+ 0 0 0 0 0 0

0 0 0 0

ADU 0 0 0 0 0 0

0 0 0 0

H4SIO 61.405 0 0 0 0 0

0 0 0 61.474

NH4NO3 0 0 0 0 0 0

0 0 0 0

LIME 0 0 0 0 0 0

0 0 0 0

CA++ 0 0 0 0 0 0

0 0 0 0

CAOH+ 0 0 0 0 0 0

0 0 0 0

CA(OH)2 0 0 0 0 0 0

0 0 0 0

CALCI(S) 0 0 0 0 0 0

0 0 0 0

AL(OH)3 0 0 0 0 0 0

0 0 0 0

FE(OH)2 0 0 0 0 0 0

0 0 0 0

MG(OH)2 0 0 0 0 0 0

0 0 0 0

Total 602056.4 0.061 1 14.123 27528.58 25.352

601795.4 0.639 27752.28 63.781

Preg-of = Pregnant leach liquor 6 = Barren liquor

Res-mu = Resin make-up Eluate = Eluate

Eluant = Eluant SiO2-out = Backwash water out

WaterSiO = Backwash water in



Table A.6: Mass balance for solvent extraction process

IN

OUT

Mass flow kg/hr

Mass flow kg/hr

ELUATE 7 A-SCRUB

LOAD 16 19

27 752.28 8 459.47 5 604.08

27 341.47 5 635.98 8 859.33

Liquids Liquids Liquids


CLINOCL 0 0 0

0 0 0

PYROPHLT 0 0 0

0 0 0

PYRITE 0 0 0

0 0 0

ALBITE 0 0 0

0 0 0

UO2 0 0 0

0 0 0

UTI2O6 0 0 0

0 0 0

UCLPO6 0 0 0

0 0 0

UO2SIO2 0 0 0

0 0 0

UO2-2W 0 0 0

0 0 0

U3O8 0 0 0

0 0 0

TIO2 0 0 0

0 0 0

K+ 0 0 0

0 0 0

AL+++ 0 0 0

0 0 0

MG++ 0 0 0

0 0 0

FE++ 0 0 0

0 0 0

FE+++ 0 0 0

0 0 0

FESO4- 82.808 0 0

74.527 8.281 0



NA+ 0 0 0

0 0 0

UO2++ 75.942 0 0

0 0 0

US2O10-- 68.141 0 0

0 0 0

US3O14-- 0 0 0

4.785 0 234.511

H3O+ 575.006 0.002 0

982.086 0.098 0.002

OH- 0 0 0

0 0 0

CL- 0 0 0

0 0 0

PO4--- 0 0 0

0 0 0

H2SO4 0 0 0

0 0 0

HSO4- 2077.817 0.04 0

0 0 0.041

SO4-- 420.579 860.464 0

2216.1 0 1025.982

HNO3 0 0 0

0 0 0

HNO2 0 0 0

0 0 0

NO3- 0 0 0

0 0 0

NO2 0 0 0

0 0 0

NO 0 0 0

0 0 0

R-NO3 0 0 0

0 0 0

R2-UC 0 0 0

0 0 0

R2-SO4 0 0 0

0 0 0

R-FEC 0 0 0

0 0 0

L4US3O14 0 0 0

0 0 0

L2-SO4 0 0 0

0 0.047 0

KEROSENE 0 0 0

0 21.581 0



ISODEC 0 0 0

0 0.048 0

NH3 0 0.009 0

0 0 0.009

AMMON-01 0 0 0

0 0 0

NH4+ 0 322.679 0

0 0 322.679

ADU 0 0 0

0 0 0

H4SIO 0 0 0

0 0 0

NH4NO3 0 0 0

0 0 0

LIME 0 0 0

0 0 0

CA++ 0 0 0

0 0 0

CAOH+ 0 0 0

0 0 0

CA(OH)2 0 0 0

0 0 0

CALCI(S) 0 0 0

0 0 0

AL(OH)3 0 0 0

0 0 0

FE(OH)2 0 0 0

0 0 0

MG(OH)2 0 0 0

0 0 0

Total 27752.28 8459.468 5604.08

27341.47 5635.981 8859.33

Eluate = Eluate Load = Solvent extraction recycle

7 = ADU recycle 16 = Spend demineralised water

A-scrub = Aqueous scrub water 19 = OK liquor



Table A.7: Mass balance for precipitation process

IN

OUT Mass flow kg/hr

Mass flow kg/hr

19 NH3 20

8 12 18 22

8 859.33 85.73 1 410.02

143.98 8 539.27 374.01 1 297.82


Liquids Liquids Liquids Solids Liquids

H2O 7276.107 64.295 1410.022

119.723 7100.497 243.085

1275.785

SIO2 0 0 0

0 0 0 0 0

MUSCV 0 0 0

0 0 0 0 0

CLINOCL 0 0 0

0 0 0 0 0

PYROPHLT 0 0 0

0 0 0 0 0

PYRITE 0 0 0

0 0 0 0 0

ALBITE 0 0 0

0 0 0 0 0

UO2 0 0 0

0 0 0 0 0

UTI2O6 0 0 0

0 0 0 0 0

UCLPO6 0 0 0

0 0 0 0 0

UO2SIO2 0 0 0

0 0 0 0 0

UO2-2W 0 0 0

0 0 0 0 0

U3O8 0 0 0

0 0 0 0 0

TIO2 0 0 0

0 0 0 0 0

K+ 0 0 0

0 0 0 0 0

AL+++ 0 0 0

0 0 0 0 0

MG++ 0 0 0

0 0 0 0 0



FE++ 0 0 0

0 0 0 0 0

FE+++ 0 0 0

0 0 0 0 0

FESO4- 0 0 0

0 0 0 0 0

NA+ 0 0 0

0 0 0 0 0

UO2++ 0 0 0

0.003 0.165 0 0 0

US2O10-- 0 0 0

0 0 0 0 0.004

US3O14-- 234.511 0 0

0 0 0 0 0

H3O+ 0.002 0 0

0 0 0 0 0

OH- 0 0 0

0 0 0 0 0

CL- 0 0 0

0 0 0 0 0

PO4--- 0 0 0

0 0 0 0 0

H2SO4 0 0 0

0 0 0 0 0

HSO4- 0.041 0 0

0 0 0 0 0.002

SO4-- 1025.982 0 0

18.738 1111.32 0.017 0 17.016

HNO3 0 0 0

0 0 0 0 0

HNO2 0 0 0

0 0 0 0 0

NO3- 0 0 0

0 0 0 0 0

NO2 0 0 0

0 0 0 0 0

NO 0 0 0

0 0 0 0 0

R-NO3 0 0 0

0 0 0 0 0

R2-UC 0 0 0

0 0 0 0 0

R2-SO4 0 0 0

0 0 0 0 0

R-FEC 0 0 0

0 0 0 0 0



L4US3O14 0 0 0

0 0 0 0 0

L2-SO4 0 0 0

0 0 0 0 0

KEROSENE 0 0 0

0 0 0 0 0

ISODEC 0 0 0

0 0 0 0 0

NH3 0.009 21.432 0

0 0 0 0 0

AMMON-01 0 0 0

0 0 0 0 0

NH4+ 322.679 0 0

5.518 327.286 0.005 0 5.012

ADU 0 0 0

0 0 0 0 0

H4SIO 0 0 0

0 0 0 130.904 0

NH4NO3 0 0 0

0 0 0 0 0

LIME 0 0 0

0 0 0 0 0

CA++ 0 0 0

0 0 0 0 0

CAOH+ 0 0 0

0 0 0 0 0

CA(OH)2 0 0 0

0 0 0 0 0

CALCI(S) 0 0 0

0 0 0 0 0

AL(OH)3 0 0 0

0 0 0 0 0

FE(OH)2 0 0 0

0 0 0 0 0

MG(OH)2 0 0 0

0 0 0 0 0

Total 8859.33 85.727 1410.022

143.983 8539.268 243.107 130.904 1297.819

19 = OK liquor 8 = ADU bleed

NH3 = Ammonia 12 = ADU recycle

20 = Wash water for centrifuge 18 = ADU product

22 = Spent liquid



A.2. Energy balance

The approach for calculating the energy balance is to consider each section individually.

Additional energy is required at the leaching and precipitation units to increase the

temperature as required. Each process where significant energy transfer or generation

occurs is seen as a black box to validate the energy balance. Finally an energy balance

over each unit in the uranium extraction plant is given. In this section it is assumed that the

ambient temperature is 20 °C.

The energy balance for the leaching pachucas is done by calculating the amount of steam

required to heat the pachuca contents to the required temperature. The feeds into the

pachuca is simplified to ore, water, air, and steam. The energy balance over the leaching

pachucas is given in Table A.8.

Table A.8: Leaching pachucas energy balance

IN

Stream Mass flow

(kg/hr) Heat capacity

(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Water 395175.0 4384.0 0.0 0.0

Ore 530000.0 800.0 0.0 0.0

Steam 59197.9 4246.0 155.0 10822313.8

Air 10000.0 1005.0 0.0 0.0

Total 10822313.84

OUT

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Water 454372.8589 4384 16 8853202.725

Ore 530000 800 16 1884444.444

Air 10000 1005 16 44666.66667

Total 10782313.84



The heat capacities for the water, steam (sub-cooled water, 175 °C, 12 bar) and air is

derived from literature (Koretsky, 2004: 507). The heat capacity for the ore is assumed to be

820 J/kg.K which is derived by taking the average of the heat capacities of several

components of the ore. The heat capacity of quartz which contributes approximately 70% of

the ore weight is 742 J/kg/K (Hemingway,1987: 275) and for muscovite which contributes

approximately 10% of the ore weight is 818 J/kg/K (Cemič,2005: 91). From this average

heat capacity is estimated at 800 J/kg.K. Mass flows for the streams are approximated from

the mass balance. The reference temperature used for the calculations in Table A.8 is

20 °C.

Further it is assumed that the air for agitation is fed at 10 000 kg/hr at a feed pressure of 6

bar. The energy loss to the environment through evaporation and conduction is assumed as

40 kW which is less than 1 % of the system energy. The pachuca content is first heated to

30 °C and thereafter the temperature of each fourth pachuca is increased with 2 °C to

maintain the desired temperature. For this reason the outlet temperature is taken as 36 °C,

since there are three pachucas that require additional heating. Using solver in Microsoft®

Office Excel, the mass flow of the steam into the system is solved to satisfy the energy

balance. The steam required for the leaching pachucas is 60 000 kg/hr.

The leach product flows to the counter-current decantation section to wash the uranium

containing liquids from the solids, recovering the product. This process takes place in large

open tanks with a high residence time. The wash water entering the counter-current

decantation from the ion exchange system is at 20 °C and therefore does not introduce

thermal energy into the system. Here the liquids and solids lose all their thermal energy to

the environment through evaporation and conduction. In the energy balance shown in Table

A.9 it is assumed that the liquids and slurry leaving the counter-current decantation section

are at 20 °C, the same as the surrounding environment and therefore all potential thermal

energy leaves the system.



Table A.9: Counter-current decantation energy balance

IN

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Ore 530000 800 10 1177778

Water 395175 4384 10 4812353

Wash water 610000 4384 0 0

Total 5990131

OUT

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Ore 530000 800 0 0

Water 403298.1 4384 0 0

Leach liquor 600000 4384 0 0

Total 0

Table A.9 shows that an energy loss of approximately 6 MW to the environment. This is

possible since a large open area will enhance evaporation which causes great energy

losses. It is also important to note that liquid surface is in continuous circular movement,

enhancing evaporation. The wash water entering the system is at 20 °C which will create an

energy gradient for energy transfer. It is calculated that the temperature of the exit streams

are 24 °C if no energy loss is assumed, but this is highly unlikely.

The slaked lime and slurry stream entering the neutralisation pachuca are at the reference

temperature (20 °C). Since neutralisation reactions occur in this process it is important to

take the energy generation caused by these reactions into account. The heat of reaction for

several of the more important reactions, such as the precipitation of metal hydroxides due to

the high pH and neutralisation reactions are calculated. From this it is calculated that

approximately 1.6 MW is produced by the reactions in the system. The energy balance for

the neutralisation process is given in Table A.10.



Table A.10: Neutralisation energy balance

IN

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Water 403298.0 4384.0 0.0 0.0

Ore 530000.0 800.0 0.0 0.0

Slaked lime 57398.0 4384.0 0.0 0.0

Solid lime 10186.0 2845.2 0.0 0.0

Reactions 1645318.0

Total 1645318.0

OUT

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Water 461439.0 4384.0 0.0 1359756.7

Ore 531047.0 800.0 0.0 285561.3

Total 1645318.0

The heat of reaction for the reactions is obtained from the Facility for the analysis of

chemical thermodynamics (2009). The reactions are the only source of energy in the system

as seen in Table A.10. If it is assumed that no energy loss occurs, and therefore leaves the

system in the exit stream. The temperature of the exit stream is solved using Microsoft®

Office Excel. Table A.10 shows the exit temperature under these conditions increases with

7 x 10-4 °C. If energy loss to the environment occurs the temperature increase is even

lower, therefore it is assumed that the temperature increase over the neutralisation process

is negligible.

No energy increases are present in the ion exchange or solvent extraction. The small

amounts of reactions occurring in these sections are highly diluted and are therefore

negligible. From this it is assumed that the temperature increase over these sections is also

negligible.



At the precipitation the reactor temperature is controlled at 30 °C to enhance reaction

kinetics. This temperature is achieved by heating the OK-liquor in a heat exchanger using

steam. An ammonia gas and air mixture is also introduced to the precipitation reactors. It is

found in the calculations, similar to that of neutralisation, that the thermal energy generated

by the reactions has no significant temperature effect. The energy balance for the

precipitation reactors is given in Table A.11.

Table A.11: Precipitation energy balance

IN

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Water 8.1 4384.0 0.0 0.0

NH3 21.4 2058.8 0.0 0.0

air 50.0 1006.0 0.0 0.0

Q

725759.9

Total 725759.9

OUT

Stream Mass flow


(J/kg.K) T-Tref

(°C) Thermal power

(Watt)

Water 8.1 4384.0 10.0 356945.3

ADU 130.9 281.8 10.0 368814.6

Air 50.0 1006.0 10.0 503080.5

Total 725759.9

In Table A.11 it is seen that 0.75 MW is required to heat the reactor contents to the desired

30 °C. All the energy transferred to the precipitation reactors leaves with the product to the

thickener, where all the energy is lost to the environment through evaporation and

conduction.

It should be noted that these calculations are only an approximation for the complex

hydrometallurgical system. From these calculations it is concluded that the energy

distribution over the uranium extraction plant follows the law of energy conservation and that

no heat integration is practical.


254 Appendix B: Equipment sizing

Appendix B: Equipment sizing

Appendix B consists of the description of the sizing methods followed for Section 3.7.

B.1. Leaching

The kinetics for the nitric acid leaching of uranium is found in Ikeda et al. (1995) as Equation

B-1

[ ] − = +

2.31 2a 1 2 2 3

H Hr A exp A exp HNO NORT RT (B-1)

For Equation B-1 the constants are given in Table B.1.

Table B.1: Leaching kinetics constants

Constant Description Value

A1 Frequency factor 2.2 x 104

A2 Frequency factor 0.46

H1 Activation energy -79 500

H2 Activation energy -36 800

The design equation used is for a batch reactor and is manipulated to give Equation B-2

(Fogler, 2006: 70):

= × ×aa

dN r S Ndt (B-2)

It was assumed that the shrinking core model is applicable to the leaching of uranium ore.

Ikeda et al. (1995) derived the shrinking core model for the reaction kinetics. Since other

side reactions occur during the leaching process, the reduction in particle radius is taken into

account with an additional linear constant. This derivation is given in Equation B-3.



−× = − + × ρ×

8ar MWdr 2.95 10dt 100 (B-3)

The extra term (2.95x10-8) in Equation B-3 is manipulated to reduce the radius of the particle

to10% of the initial radius. Further it is assumed that the uranium content of the particles is

0.034%. The amount of particles (N) is calculated using Equation B-4.

=Total solid feed volumeN

Volume of single spherical particle (B-4)

The feed nitrate concentration ([NO3-]0) is taken from the mass balance in Appendix A. It

was derived from the reaction stoichiometry that the rate at which the nitrates decrease in

the system is given by Equation B-5.

− − = − × a

3 3 00

N2NO NO3 v (B-5)

In Equation B-5 v0 is the volumetric flow of the liquid feed. The above equations were

implemented in Polymath® to determine the influence of several design parameters. The

Polymath® program is shown next.

A1 = 2.2e4

A2 = 0.46

R = 8.314

T = 273 + 30

MW = 270

rho = 2.650

GramU0 = 530 * 0.34 * 1000 / 60

MolU0 = GramU0 / MW

CHNO2 = 0

C0NO3 = 0.0618

ra = (A1 * exp(-79500 / R / T) + A2 * exp(-36800 / R / T) * CHNO2) * (CNO3) ^ 2.3

w = 1 / ra

d(r)/d(t) = -(ra * MW / rho / 100 + 0.0000000295)

r(0) = 0.0000375



S = 4 * (22 / 7) * (r ^ 2) * (100 ^ 2)

d(MolU)/d(t) = ra * S * N

MolU(0) = 0

GramU = MolU * MW

dXdt = ra * S * N / MolU0

N = 6.15e17 * (0.34 / 1000)

CNO3 = C0NO3 - (2 / 3) * (MolU / 381645)

t(0) = 0

t(f) = 1137

X = GramU / GramU0

Using the above code the influence of temperature and nitrate concentration was determined

in order to optimize the system. Firstly the influence of the nitrate concentration was

determined, while a constant temperature was used. It is important to note that the

concentration of HNO2 was assumed to be zero, since it is not fed as raw material and it has

no effect on the overall chemical reaction (given in Equation B-6).

(B-6)

The concentration of the nitrates in the feed was varied between 0.06 and 0.1 mole/L, at a

temperature of 25 °C. The influence of the nitrate concentration is shown in Figure B.1.

− + ++ + → + +22 3 3 2 23UO 2NO 8H O 3UO 12H O 2NO



Figure B.1: Influence of nitrate concentration on reaction kinetics

In Figure B.1 it is seen that an increase in nitrate concentration results in a significant

increased amount of product. It is also noticed, from Figure B.1, that the amount of uranium

leached increases exponentially as the nitrate concentration increases. This effect is caused

by the exponential function of the reaction kinetics in respect to the nitrate concentration.

The influence of temperature was found to have a significant effect on the reaction kinetics.

The temperature was varied between 25 and 55 °C while keeping the nitrate concentration

constant at 0.0182 mole/L. Figure B.2 illustrates the results of this variation of temperature

on the reaction kinetics.

0

2

4

6

8

10

12

14

16

18

20

0 200 400 600 800 1000 1200

Mol

e ur

aniu

m le

ache

d

Time (min)

0.06 mole/L 0.07 mole/L 0.08 mole/L 0.09 mole/L 0.1 mole/L



Figure B.2: Influence of temperature on reaction kinetics

As seen in Figure B.2, the temperature also has a significant effect on the reaction kinetics.

The influence of the temperature forces the use of steam to ensure a proper conversion in

the leaching section, especially during the winter. It is therefore important to optimise both

the temperature and the nitrate concentration to ensure economical conversion.

Taking the cost considerations for temperature increase and nitrate addition into account, it

was optimised and found that the nitrate recycle from solvent extraction is sufficient while the

temperature is controlled at 30 °C. It was also approximated that the heat loss will cause a

temperature drop of about 0.55 °C per tank and therefore it was decided to heat each fourth

leaching tank.

The sizing of the leaching process is done according to the graphical method described by

the Minerals Council of Australia (2006). The kinetic data is obtained at a nitrate

concentration of 0.0618 mole/L and a temperature of 30 °C. Figure B.3 illustrates the

conversion rate against the conversion of uranium, and from it the existing stages are

evaluated.

0

1

2

3

4

5

6

7

8

0 200 400 600 800 1000 1200

Mol

e ur

aniu

m le

ache

d

Time (min)

55 °C 45 °C 25 °C 35 °C



Figure B.3: Leach tank size evaluation

From Figure B.3 it was found that 11 of the existing leaching stages will give a uranium

conversion of 96%. From this it is concluded that 11 of the existing leaching tanks will be

used while maintaining a temperature of 30 °C and a nitrate feed concentration of 0.0618

mole/L.



B.2. Counter-current decantation

The existing counter-current decantation equipment consists of six thickeners in series. The

first thickener has a larger area to ensure minimal solids in the feed to the ion exchange

section. For the expansion project a larger feed must be processed in this section. It was

decided to split the feed into two trains, Train 1 consisting of the existing thickeners while

Train 2 will consist of six new thickeners in series. The smaller capacity of Train 2 is sized

according to the existing equipment capacity and some guidelines from the literature.

The capacity of the existing thickeners is calculated from available plant data. Thickener 1

of the existing train has a cross-sectional area of 2 800 m2 and that of Thickener 2 to 6 is 2

400 m2. Currently Train 1 is processing a leach liquor from an ore feed of 8 000 ton/day

from which a capacity coefficient is calculated by dividing the cross-sectional area with the

ore feed. This capacity coefficient for Thickener 1 is 0.35 and for Thickener 2 to 6 it is 0.3

m2 per ton per day. This correlates with figures from Merritt (1971) which suggest this

capacity coefficient should be between 0.23 and 0.6 m2 per ton per day. The cross-sectional

area of the thickeners of Train 2, which should process an ore feed of 5 000 ton/day, is sized

directly from this and was found to be 1 750 m2 for Thickener 1 and 1 500 m2 for Thickener 2

to 6. From this the diameter of Thickener 1 of Train 2 is approximated as 50 m and for

Thickener 2 to 6 as 45 m. The total capacity of the two trains will process an ore feed of 13

000 ton/day.

The wash water that enter Thickener 1 in both Train 1 and 2 is calculated with Equation B-7.

= n

100%Uranium loss(wash ratio) (B-7)

In Equation B-7, n is the number of stages in the counter current decantation section, which

is 6 in this case. A uranium loss of 0.01% is assumed resulting in a wash ratio of 1.

Therefore the volumetric flow of the wash solution is equal to that of the leach liquor.



B.3. Ion exchange

For the ion exchange section fixed-bed column ion exchange process is used. The sizing of

the columns was done using resin specifications and basic mass balance calculations. The

Ambersep 400 resin is used which has a capacity of 1.4 eq/L and according to Merritt (1971)

40% of the resin capacity is occupied by uranyl complexes. The bed voidage is assumed to

be 30%.

From the kinetics and resin capacity either the column dimensions or the adsorption time

can be calculated. The column dimensions of the existing adsorption columns are used. It

was calculated that the resin capacity is 78 g UO2(SO4)34- per litre resin and from the mass

balance the uranium content is 0.339 g UO2(SO4)34- per litre solution. A mass balance over

this system shows that the adsorption time for one column is 20 hours. Therefore during this

period there is sufficient time for elution and washing of the resin.

The elution of saturated resin is done with seven bed volumes of eluant (acidified NH4NO3).

This amount of eluant is chosen to ensure the desired concentration of uranyl complexes

(4 g U3O8/L) in the feed to the solvent extraction. According to Merritt (1971) the residence

time for eluant in the column should be between 12 and 20 minutes. A residence time for 20

minutes was assumed for this design which resulted in an elution time of 8 hours. From this

it is seen that the elution is easily done in the available time with ample time left for washing.

A total volume of 500 m3 of eluate is sent to the solvent extraction feed tank every 20 hours,

resulting in a volumetric flow of 25 m3/hr with a uranium content of 3.9 g U3O8/L.

B.4. Precipitation

The precipitation of ADU is achieved by bubbling ammonia and air mixture through the OK

liquor. The existing reactor has a diameter of 1.5 m and height of 1.5 m with a volume of

2.65 m3. The tank is operated at 60% of its capacity to ensure adequate space for gasses

and emissions (Motsau, 2008:64). Therefore the effective volume of the reactor is 1.6 m3.

The reaction kinetics as described by Motsau (2008) is given in Equation B-8.

− = 2A Ar kC

(B-8)



In Equation A-7 –rA is the rate at which ammonia reacts (mole/m3.min), k is the reaction rate

constant with a value of 857.8 m3/mole.min and CA is the concentration of UO2(SO4)34- in the

solution. The design equation for a continuous stirred tank reactor (CSTR) is taken from

Fogler (2006) and shown in Equation B-9

=−

A02

A0

vC XVk(C (1 X)) (B-9)

In Equation B-9 X is the conversion of UO2(SO4)34- to ADU, V is the effective volume of the

reactor, CA0 is the feed concentration and v is the volumetric flow of the OK liquor.

From the available plant data for the existing precipitation reactor and the OK liquor

concentrations and flow, the conversion was calculated. Currently the concentration of

UO2(SO4)34- is 39 mole/m3 and the volumetric flow rate is 6 m3/hr for the OK liquor. This

gives a conversion of 99.86%. The same reactor is evaluated to determine whether its

capacity is sufficient for the increased OK liquor flow. From the mass balance the increased

OK liquor has a volumetric flow rate of 7.8 m3/hr and UO2(SO4)34- concentration of 44

mole/m3. The conversion obtain from the increased OK liquor is 99.85% which is a very

small decrease in conversion. Therefore the existing precipitation reactor will be used to

process the increased OK liquor and will result in 173 mole ADU/hr which is sent to a

thickener.

The existing thickener has a diameter of 15 m and a height of 4 m, giving a cross-sectional

area of 180 m2. Currently this thickener processes 1.8 ton ADU/day and from this the

capacity coefficient is calculated as 98.17 m2 per ton per day. According to Weiss (1985)

the capacity coefficient should be between 4.7 and 11.6 m2 per ton per day but the

calculated coefficient is much higher which shows the thickener is over-designed. For the

increased feed of 2.6 ton ADU/day the capacity coefficient is 68 m2 per ton per day which is

still higher than the prescribed range. Therefore the existing ADU thickener is used to

process the increased feed.



B.5. Neutralisation

The pH of the slurry from the counter current decantation section must be 10.5 before it is

sent to the gold extraction plant. Therefore a neutralisation reactor is necessary to

neutralise this slurry by adding lime. Since only 11 of the 14 existing leaching pachucas are

used, one of the remaining pachucas is used to neutralise the slurry. This pachuca are

already equipped with air agitation to ensure good mixing of the slurry and lime before it is

sent to the gold extraction plant.

From basic calculations involving the solubility constants of several metal hydroxide species,

the mass of lime required to reach the desire pH was found as 10 700 kg/hr. This lime is

slaked with 46 700 kg/hr water to achieve a slaked lime feed density of 1.15 kg/L. The total

volumetric flow to the neutralisation pachuca , including the slaked lime and slurry feed, is 60

m3/hr. For the effective volume of 750 m3 for the neutralisation pachuca, a residence time of

1 hour and 15 minutes is achieved, which is sufficient to neutralise the slurry.


264 Appendix C: Detail design calculations

Appendix C: Detail design calculations

Appendix C consists of a detailed explanation for the calculations and logical reasoning

followed during the detail design. For this feasibility study, the detail design is done for the

solvent extraction section.

C.1. Loading isotherms

The loading isotherm for the extraction section is obtained using Equation 4-1 and the given

constants. The loading isotherm using the reported constant does not fit the experimental

data; therefore the constants are manipulated to obtain a satisfactory modelled isotherm. In

Figure C.1 the experimental data is modelled using Equation 4-1 with the respective

constants.

Figure C.1: Modelling of the extraction loading isotherm

0

1

2

3

4

5

6

0 0.5 1 1.5 2 2.5 3 3.5 4 4.5 5

y (U

org)

(g/L

)

x (Uaq) (g/L)Model - repoted constants Model - manipulated constants Experimental data



In Figure C.1 it is noticed that the reported constants for Equation 4-1 does not fit the

experimental data well. The manipulated constants however fit the experimental data well

and are used to obtain the loading isotherm for the specific operating conditions.

The experimental data obtained from Figure 4.2 is shown in Figure C.2. The data is plotted

in Excel to obtain a smooth line between the data points.

Figure C.2: Loading isotherm data for the stripping section

The operating line is calculated from a basic mass balance of the uranium over the process,

given in Figure C.3.

Figure C.3: Solvent extraction process flow schematic

0

5

10

15

20

25

30

35

40

0 1 2 3 4 5 6 7 8 9

y (U

3Oaq

) (g/

L)

x (U3O8org) (g/L)



In Figure C.3, x represents the uranium content (g U3O8/L) of the phase the uranium is

transferred from while y represents the uranium content of the phase the uranium is

transferred to. The operating lines for the extraction and stripping sections are derived from

the mass balance and given in Equation C-1 and C-2 respectively.

= − +A AE,in E,out E,in E,out

O O

V Vy y x xV V (C-1)

= − +O OS,in S,out S,in S,out

A A

V Vy y x xV V (C-2)

As seen in Figure C.3, the extraction and stripping section are interdependent due to the

recycling of the organic solvent. The uranium content of the aqueous stream, xE,in, into the

extraction section and the aqueous stream, yS,out, out of the stripping section is specified for

the design. A recovery of 0.98 is assumed in the extraction section, which sets the value for

yE,out. The uranium content of the loaded solvent from the extraction section, yE,out, is equal

to that of the solvent fed to the stripping section, xS,in. The process is iteratively solved to

obtain equal value for yE,in and xS,out. This is achieved by varying the ratio of the volumetric

flows of the aqueous, VA, and the organic, VO, phases for the stripping section. The VA to VO

ratio for the extraction section is specified as 1.1:1.

The resulting McCabe-Thiele graphs from the procedure discussed above are shown in

Figures C.4 and C.5. These graphs are the final result in the iterative solving, involving four

iterations, to obtain equal values for yE,in and xS,out.



= +E,inE,out E,in

O

A

0.98xy yV

V

Figure C.4: McCabe-Thiele for the Extraction section

The calculated operating line for the extraction section, seen in Figure C.4, is given in

Equation C-3.

= − × +E,in E,out1 1y 3.5918 3.751 x

1.1 1.1 (C-3)

The operating line for the extraction section is fixed due to the specified VA to VO ratio. For

the first iteration of McCabe-Thiele the value for yE,in was assumed to be 0 g U3O8/L from

which a new value of 0.25 g U3O8/L for yE,in is found. The value for yE,out in Equation B-1 is

calculated from the assumed recovery of 0.98 in the extraction section (see Equation C-4).

This gave the final operating line given in Equation C-3.

(C-4)



= − +O OS,in S,out

A A

V Vy 12 3.5918 xV V

The efficiency used in the McCabe-Thiele method is taken as 0.98, which depends on the

level of mixing and residence time in the mixers. The value for yE,out is very close to the

assumed value of 0.25 g U3O8/L, which is satisfactory due to the human factor in the

McCabe-Thiele method. Once the McCabe-Thiele method gives satisfactory results, the

method is applied to the stripping section to iteratively solve the mass balance as mentioned

above.

Figure C.5: McCabe-Thiele for the Stripping section

The calculated operating line for the stripping section, seen in Figure C.5, is given in

Equation C-5.

(C-5)



The value for yS,out is specified as 12 g U3O8/L in Equation B-5, while the value for xS,in is

equal to yE,out (3.5918 g U3O8/L). The value for the VO to VA ratio is varied to obtain a value

of 0.25 g U3O8/L for xS,out. The final ratio for VO to VA is obtained as 3.6 which then solves

the mass balance as seen in Figure C.4. From the McCabe-Thiele method the uranium

loading (g U3O8/L) exiting each stage in the extraction and stripping sections are obtained.

C.2 Mixer design

The residence time is used to calculate the volume of the liquids in the mixer box. An over-

design of 10% was used throughout the calculations for the mixer design. In Equation C-6

the volume of the liquids (V in m3) was calculated by multiplying the residence time (t in hr)

with the aqueous and organic volumetric flow rate (QA and QO) and the over-design. The

calculation is shown in Equation C-6.

= × + ×A OV t (Q Q ) 1.1 (C-6)

The residence time is specified as 2 minutes to ensure height stage efficiency and the total

design volumetric flow (QA + QO) is 54.2 m3/hr. The liquid volume is calculated as 2 m3. A

rule of thumb is used to determine the dimensions of the mixer box which states that the

height is equal to the diameter of the mixer. Using the above information the dimensions of

the mixer box is as follows:

• A height of 1.7 m, which includes an extra 0.3 m for instrumentation.

• A diameter of 1.4 m.

The diameter of the impeller blade is approximated as 80% of the mixer box diameter and

results in a 1.1 m impeller diameter. The used impeller acts as a mixer and a pump and is

supplied by MC Process revered to as the MC Process SX impeller (Watson, 2009). The

three mixer design equations and their constant, for the specific impeller, as well as the

equation for the hydraulic efficiency follows.



( )+= =

=

= =π

πε =

A Oq

p

p

h 2 2 2p

2h q

p

Q Q 1.1N 0.01048

ND

N 0.5493

2gHN 2.46N D

N N2N

(C-7)

(C-8)

(C-9)

(C-10)

The most important parameters in the above equations are N and Dp, which is the impeller

rotary speed (revolutions per second) and the impeller diameter (m) respectively. Due to the

fact that the impeller acts as both a mixer and pump, it is also important to calculate the

head produced H in meters, and the hydraulic efficiency, ε.

Since the over-designed volumetric flow and the impeller diameter is known, it is possible to

calculate the impeller rotary speed using Equation C-7. The tip speed of the impeller is a

crucial design parameter, which determines the level of crud formation. Crud is an over-

emulsified mixture of organic and aqueous phases which does not settle. An approved

design prescribed by MC Process has a tip speed of 3.9 m/s, compared to the 4 m/s

resulting from the above design. It is therefore accepted that the designed impeller rotary

speed is within the design bounds. Once the impeller rotary speed is known, the produced

head and hydraulic efficiency is calculated using Equation C-9 and C-10.

The produced head should be able to supply sufficient pressure increase to ensure

continuous flow into the settler vessel. According to Eckhart (2004) stirring devices always

work in the turbulent flow regime, thus ReR (Equation C-7) should be larger than 5x105.

Using Equation C-11 together with the calculated impeller rotary speed the Reynolds

number, ReR, is calculated to indicate the flow regime (Eckhart, 2004:15):

=υ

2m

Rc

NDRe (C-11)

In Equation C-11 the subscript m is an indication of the agitator and c is for the continuous

phase. The symbols nR and dR is the rotor speed and mixer diameter, respectively. The

kinematic viscosity of the continuous phase is given by the symbol νc. For systems with an

aqueous-organic ratio of close to one, the continuous phase is seen as the denser phase

which in this system is seen as the aqueous phase for this system (Minerals Council of



Australia, 2006). The Reynolds number is calculated as 1.7 x 106 and is therefore in the

turbulent flow regime.

C.3. Settler design

The settler are design according to the settling kinetics, Equation 4-2 and 4-3, and constants

reported by Stönner and Wiesner (1982). The above mentioned equations are for a batch

system, which are manipulated in order to represent a plug flow reactor system. Equation C-

12 and C-13 are used in this design.

= −1 1

T

dQ Qk.H.w.dL Q (C-12)

= −

2 1 2

T T

dQ Q Qw k.H. c.dL Q Q (C-13)

In Equations C-12 and C-13, Q1 is the volumetric flow (m3/hr) of the small droplets, Q2 is the

volumetric flow of the large droplets, w is the width of the settler vessel (m) and L is the

length of the vessel (m). In the above equations QT is the total volumetric flow of the

dispersion layer, which is the sum of the volumetric flow of the aqueous (QA) and organic

phases (QO). The base of these differential equations is the length, L (m), of the vessel.

The height of the dispersion layer, H, is assumed to be directly proportional to the total

volumetric flow of the droplets. The initial height of the dispersion layer is dependent on the

w, QT and the velocity of the dispersion through the vessel (v in m/hr), see Equation C-14.

= T0

QHw.v (C-14)

In Equation C-14, H0 is the initial height of the dispersion layer and QT is the total volumetric

flow (aqueous and organic) into the settler vessel. The values of the constants k and c

(settling velocity of large droplets) in equation C-12 and C-13 is reported as 17.6 hr-1 and

13.2 m/hr respectively. In the design of the extraction settlers, the droplets contain the

organic solvent, while the aqueous phase is the continuous phase. It is further assumed that

the total volumetric flow of droplets is initially small droplets. This assumption is valid if the



mixing efficiency is high and will result in the slowest possible settling rates. The slowest

settling rates are chosen to act as an over-design for the settler vessel.

Combining the above equations, a Polymath® program is created to simultaneously solve the

two differential equations, which follows. This program solves the differential equations (C-

12 and C-13) in terms of the length. The vessel is designed with an adequate settling length

to allow the height of the dispersion layer, H, to reach a desired height.

d(Q1)/d(s) = -rate1 * dikte d(Q2)/d(s) = (rate1 - rate2) * dikte rate1 = k * H * Q1 / QT rate2 = c * Q2 / QT Q1(0) = 28.317 Q2(0) = 0 s(0) = 0 s(f) = 5.5 speed = 25 dikte = 2 k = 17.6 c = 13.2 Q0 = 28.317 QT = (Q1 + Q2) + (Q1 + Q2) / 1.1 H0 = (Q0 + Q0 / 1.1) / (dikte * speed) H = H0 * QT / (Q0 + Q0 / 1.1)

In the above Polymath® program there are two variable design parameters (w and v). These

parameters are varied to determine their influence on the settling kinetics. The extraction

and stripping sections differs in regard to the VA to VO ratio, therefore it is sized separately.

The extraction section is sized first where a VO to VA ratio of 1.1:1 is used with the aqueous

feed of 25.8 m3/hr. The width parameter, w, is varied between 2 and 4 m, while keeping v

constant at 30m/hr and is shown in Figure C.6.



FigureC.6: Vessel width sensitivity analysis

In Figure C.6 it is seen that the vessel width has a significant influence on the initial height of

the dispersion layer, while having an insignificant influence on the required length of the

vessel. Figure C.6 shows that the initial height of the dispersion layer decreases when the

vessel width increase. In Figure C.7 the velocity (v) is varied between 30 and 55 m/hr, while

keeping the width at 1.5 m.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 3 6 9 12 15

Hei

ght -

H (m

)

Length -L (m)

2 m

3 m

4 m



Figure C.7: Dispersion layer velocity sensitivity analysis

From Figure C.7 it is noticed that the dispersion layer velocity has a significant effect on its

initial height, while influencing the required length to some extent. Both the vessel width and

the dispersion layer velocity effects the dimensions of the required settler vessels, which

directly influence the equipment costs. Therefore it is important to do a sensitivity analysis

combining the vessel width and dispersion layer velocity to investigate the economical

influence. This economical sensitivity analysis is based on the area of equipment material

and is shown in Figure C.8.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 3 6 9 12 15

Hei

ght -

H (m

)

Length - L (m)

30 m/hr

35 m/hr

40 m/hr

45 m/hr

50 m/hr

55 m/hr



Figure C.8: Economical sensitivity analysis on the extraction settlers

The cost of equipment is directly proportional to the area of materials needed for

construction. An increase in material area will increase the manufacturing and installation

cost. It is therefore important to determine the optimum configuration between dispersion

velocity of the layer and the width of the vessel. It is also important to ensure that the initial

height of the dispersion layer and the width of the vessel is within realistic bounds. A

realistic initial height of the dispersion layer is chosen as 1.5 m, which will include a 30%

over-design to ensure enough space for air. The width of the settler vessel should not be

greater than 2 m to ensure effective distribution of the dispersion at the top of the mixer. The

length calculated will ensure a final dispersion layer height of 2.5 cm. The calculated length

represents the distance from the distributor to the aqueous weir.

According to these limitations and Figure C.8 the best economical and efficient settler choice

has a width of 2 m and a dispersion layer velocity of 25 m/hr, a height of 1.5 m and a length

of 7.3 m which includes additional length of 1.8 m for the aqueous weir. The effective

settling length for the extraction settler is 5.5 m, which will ensure effective separation of the

two phases to reduce the height of the dispersion layer to less than 2.5 cm and results in a

residence time of 12.5 minutes. The final height of the dispersion layer will increase slightly

in practice and therefore the settlers are designed to reach a small final dispersion height.

0

10

20

30

40

50

60

70

0 10 20 30 40 50 60

Are

a (m

2 )

Velocity flow (m/hr)

1.5 m

2 m

3 m



The dispersion height profile over the length of the settler vessel for the specified settler

dimensions is shown in Figure C.9.

Figure C.9: Dispersion height profile over the length of the extraction vessel

It is important to note, from Figure C.9, that in the last third of the settling vessel length, the

dispersion layer height does not show a drastic decrease. This effect is expected from the

kinetic equation for the large droplet settling (Equation C-9), where the settling rate is directly

proportional to the dispersion layer height. A larger final dispersion layer height will increase

the probability of solvent loss. Due to the great economical strain associated with solvent

loss, it is decided to increase the capital cost which will decrease the operating cost of the

solvent extraction section. The capital cost is increased by increasing the vessel length

which will reduce the solvent loss, therefore, decreasing the operating cost.

The settler vessels for the stripping section are sized next. The same influence of the

dispersion layer velocity and the vessel width, as seen in Figure C.6 and C.7 for the

extraction section, is observed for the settling efficiency for the stripping section. It is

imperative to compare the extraction and stripping kinetics to determine the compatibility. A

velocity of 25 m/hr is used for the dispersion layer and a vessel width of 2 m is used to

compare the settling kinetics. The height profile of the dispersion layer for the stripping and

extraction section is shown in Figure C.10.

0

0.3

0.6

0.9

1.2

0 0.5 1 1.5 2 2.5 3 3.5 4 4.5 5 5.5

Hei

ght -

H (m

)

Length - L (m)



Figure C.10: Economical sensitivity analysis on the stripping settlers

In Figure C.10 it is noticed that for the same dispersion layer residence time, the settling

kinetics of the stripping section is faster and will therefore require a smaller settling vessel.

However, since the material of construction is vacuum infused fibreglass, a mould is used to

construct the settling vessels. The construction cost for this manufacturing method consist

of the price for each mould and construction materials used, therefore, using one standard

vessel size will optimize construction cost. Due to the smaller vessels required for the

stripping section, it is decided that the size for both the stripping and scrubbing settler

vessels are exactly the same as for the extraction settling vessels.

According to Minerals Council of Australia (2006), it is important to consider an internal

recycle stream of the continuous phase to ensure the desired phase continuity if the feed

ratio is close to one. The internal recycle stream will result in a different internal VO to VA

ratio and a larger total volumetric flow. Figure C.11 illustrates the different height profiles as

a result of internal recycling for the extraction section.

0

0.3

0.6

0.9

1.2

0 0.5 1 1.5 2 2.5 3 3.5 4 4.5 5 5.5

Hei

ght -

H (m

)

Length - L (m)

Extraction

Stripping



Figure C.11: Height profiles for internal recycle and without internal recycle

It is seen, in Figure C.11, that the initial dispersion height increases to approximately 1.4 m

with the addition of the internal recycling stream. The height of the settler vessels is chosen

as 1.5m and is therefore adequate to handle the internal recycle of 33% of the aqueous

phase. The final dispersion layer height resulting from the internal recycle is 3 cm which is

still acceptable. Due to the faster settling kinetics of the stripping section, it is derived that

an internal recycle in the stripping section is also viable.

C.4. Pipe sizing

Due to safety considerations with regards to fire hazards, it is important to design the inside

diameter of the pipe to ensure a velocity flow of less than 1 m/s. It is also imperative to use

normal pipe size standards to reduce the material cost of construction. Equation C-15 is

used to determine the velocity flow.

= π

2Q *1.1v

ID2

(C-15)

The inside diameter (ID in meters) is varied using the standard pipe sizes to obtain a velocity

flow (v in m/s) of less than 1 m/s. The results are shown in Chapter 4.

0

0.3

0.6

0.9

1.2

1.5

0 0.5 1 1.5 2 2.5 3 3.5 4 4.5 5 5.5

Hei

ght -

H (m

)

Length - L (m)

Without internal recycle With 33% internal recycle



C.5. Pump sizing

There are two important parameters to determine when sizing a pump, the first parameter is

the head required to pump a liquid to the required position and the second parameter is the

volumetric flow of the liquid. The first parameter is calculated using the law of conservation

of energy and the second parameter is calculated in the mass balance using the law of

conservation of mass.

The head is calculated using the momentum balance which is a mathematical representation

of the law of conservation of energy. The overall momentum balance is re-written to

represent the pressure change for each type of energy and is shown in Equation C-16

(Neomagus, 2008: 34):

(C-16)

The first term, ΔPEP, in the momentum balance equation represent the change in energy due

to change in pressure. The change in pressure is equal to zero because the liquid level of

the tank and the liquid level of the settling vessel are both at atmospheric pressure. The

change in mechanical pressure due to the change in pressure is equal zero Pa. The

calculation for ΔPEP is shown in Equation C-17 (Neomagus, 2008: 33):

(C-17)

The second term, ΔPEL, in the momentum balance represents the change in energy due to

change in height, thus the change in potential energy. Since the liquid level in the tank and

the liquid level of the settling vessel are at different heights, a change in potential energy will

occur. The calculation for ΔPEL is shown in Equation C-18 (Neomagus, 2008: 33):

(C-18)

In Equation B-18 the variables are as follows: ρ = density of water (kg/m3)

g = gravitational force (m/s2)

z1 = height at point 1 with datum as

reference (m/s)

z2 = height at point 2 with datum as

reference (m/s)

∆ + ∆ + ∆ + ∆ = ∆EP EL KE f AP P P P' P

∆ = −EP 1 2P P P

∆ = ρ −EL 2 1P g(z z )



The third term, ΔPKE, in the momentum balance represents the change in energy due to

change in velocity, thus the change in kinetic energy. The velocity, of the liquids at both the

tank and settler vessel liquid level, is close to 0 m/s and will not result in a pressure increase.

The change in mechanical pressure due to the change in kinetic energy is zero Pa. The

calculation for ΔPKE is as follow (Neomagus, 2008: 33):

(C-19)

In Equation C-19 the variables are as follows: ρ = density of water (kg/m3)

ν1 = velocity at point 1 (m/s)

v2 = velocity at point 2 (m/s)

The fourth term, ΔP’f, in the momentum balance represents the change in energy due to

friction loss. There will always be friction loss due to transport of the liquid. The flow regime

is important in the calculation of the friction coefficient (f’). The equation used to calculate

the friction loss is independent on the flow regime but dependent on the roughness of the

pipe (ε), diameter of the pipe (D) and the Reynolds number of the flow (Re). The calculation

for ΔP’f is as follow (Neomagus, 2008: 41-77):

(C-20)

In Equation C-20 the variables are as follows: K = resistance coefficient

L = length of the pipe (m)

D = diameter of the pipe (m)

∆ = ρ ν − νKE 2 1P ( )

( )

( )

ρ υ=

µ

= + +

ε = − +

=

=

= Σ

+ ρν∆ =

112 12

1.5

160.9

16

T2

Tf

DRe

8 1f ' 8Re A B

7A 2.457ln 0.27Re D

37530BRe

f 'LfD

K K

f KP

2



The term on the right hand side of Equation C-16, PA, represents the amount of pressure

required from an external source to pump the liquid to the required position. If the term has

a negative value, no pump is required and gravity flow can be used. The required pressure

is translated to head (H in meters) using Equation C-21:

(C-21)

It is also important to note that the mixer-pumper used for mixing in the solvent extraction

section gives an additional head of 2 m and should be subtracted from the head calculated

using Equation C-16 an C-21. The head and volumetric flow calculated is used to determine

the purchased cost in Chapter 5.

C.6. Control valve

The control valves are design according to the rules of thumb supplied by Svreck et al.

(2008). There are two important design parameters used to size a control valve. The first is

the pressure drop across the valve (∆PCV) and is calculated using Equation C-22 (Svreck et

al.,2008: 40).

(C-22)

The pressure across the valve is ∆PCV (kPa), Ps is the supply pressure, Qm is the maximum

feed pressure, Qd is the design feed pressure and ∆Pf is the pressure loss due to friction in

the pipes. The minimum pressure drop (∆Pb) is assumed to be 10 kPa. ∆Pf is calculated

using the same methodology as discussed for the pump sizing. Once the pressure across

the control valve is calculated it is possible to calculate the valve coefficient.

The valve coefficient is defined in Svreck et al. (2008) as the numbers of US gallons that will

pass through a control valve in 1 minute, when the pressure drop across the valve is 1 psi.

From the definition it is seen that the units is US metric, however, there exists a conversion

factor for this problem. The equation used to calculate the valve coefficient is shown in

Equation C-23 and the conversion factor is shown in Equation C-24 (Svreck et al., 2008:41).

(C-23)

∆=

ρAPH

g

∆ = + − ∆ + ∆

m

CV s f bd

QP 0.05P 1.1 1 P PQ

=∆VQC

PSG



(C-24)

Equation C-23 is used using SI units and the CV is calculated and converted to determine

the CV as per definition. Table C.1 illustrates the calculated variables for Equation C-22 and

C-23.

Table C.1: Variables for Equation C-22 and C-23.

Qm (m3/hr) Qd (m3/hr) ∆Ps (kPa) ∆Pf (kPa) ∆Pb (kPa) SG

U03-FCE01 33.00 30.00 95 12 10 1

U03-FCE02 28.40 25.82 134 12 10 1

U03-FCE03 28.40 25.82 134 12 10 1

U03-FCE04 33.00 30.00 95 12 10 1

U03-FCE05 31.24 28.40 129 12 10 0.817

U03-FCE06 6.25 5.68 138 12 10 1

U03-FCE07 11.00 10.00 138 12 10 1

U03-FCE08 11.00 10.00 120 12 10 1

U03-FCE09 5.50 5.00 120 12 10 1

U03-FCE10 8.68 7.89 138 12 10 1

U03-FCE11 33.00 30.00 95 12 10 1

U03-FCE12 2.20 2.00 138 12 10 1

U03-FCE13 2.20 2.00 138 12 10 1

U03-FCE14 2.20 2.00 138 12 10 1

U03-FCE15 8.68 7.89 119 12 10 1

U03-FCE16 33.00 30.00 95 12 10 1

U03-FCE17 5.50 5.00 95 12 10 1

U03-FCE18 11.00 10.00 95 12 10 1

U03-FCE19 31.24 28.40 138 12 10 1

U03-FCE20 31.24 28.40 95 12 10 0.817

Combining the variables in Table C.1 and Equation C-23 and C-24, the CV is calculated and

shown in Table 4.6. The calculated CV is used to determine which valve to use once the

control valve supplier is contracted.

=3

0.5 0.5

gpm m1 0.865psi hr.bar


283 Appendix D: Techno-economic calculations

Appendix D: Techno-economic evaluation

D.1. Capital investment

The capital investment is calculated using the delivered-equipment cost method according to

Peters et al. (2004). For this method the equipment purchased cost needs to be calculated for

all the new equipment used. The Marshall & Swift (M&S) cost index is used to calculate most of

the equipment purchased cost. Cost indexes are used due to inflation in the value of money.

The M&S index values are published in each issue of the Chemical Engineering magazine, and

the starting value for the M&S index in 1926 equals to 100 and the present value equals 1504.8

(Lozowski, 2009: 56).

The time value of money can be corrected by adjusting for inflation/deflation. This adjustment

can be done by the M&S index values using Equation D-1 (Ulrich & Vasudevan, 2009: 49).

=

s

P,v,s P,v,rr

M& SC CM& S (D-1)

In Equation D-1 the subscripts r and s is the different years in which the equipment is

purchased, with purchased cost CP,v,i. Equation D-1 should be used for all purchased cost

calculated from a cost graph instead of a mathematical calculation. The formulas that is used

for these calculations is given and in Table D.1 a summary of these costs are given

Mixer-settler The mixer-settlers chosen for this process is specialized equipment which means that

conventional cost estimations with calculations involving M&S indexes will give unrealistic

purchased cost for the equipment. The company providing the equipment, MC Process, gave a

ball park figure for the process as R 14 000 000. This estimation includes the agitators,

impellers, mix boxes, settlers, instrumentation, fire protection, and design time (Watson, 2009).



Heat exchanger The heat exchanger purchased cost is calculated using Equation D-2 (Douglas, 1988: 572).

( ) =

0.65c

M& SPurchased cost $ 101.3A F280 (D-2)

In Equation D-2, the symbols are: A = area (ft2)

Fc = (Fd+Fp)Fm

Fm = shell and tube material

Fd & Fp = correction factors for heat exchangers

The values for the correction factors are supplied in Douglas (1988).

Pressure vessels / storage vessels

The purchased cost for pressure vessels / storage vessels is calculated using Equation D-3

(Douglas, 1988; 574).

( ) ( ) =

1.066 0.82c

M& SPurchased cost $ 101.9D H F280 (D-3)

From Equation D-3 the symbols are describes as: D = diameter (ft)

H = height (ft)

Fc = Fm x Fp

Fm & Fp = correction factors for vessel

material and pressure

The values for the correction factors are supplied in Douglas (1988).

Pumps The purchased cost of the pumps is calculated using Figures D.1 and D.2. Figure D.1 is the

purchased cost of the pump including pump, steel base, and coupling, but no motor. Figure D.1

also provides correction factors for material of construction and provides the pump power

requirements in kilowatts (Peters et al., 2004: 517). Figure D.2 is the purchased cost of the

electric motor needed for the pump (Peters et al., 2004: 520).



Figure D.1: Cost of general-purpose centrifugal pumps

Figure D.2: Cost of electric motors



Thickeners

The installed cost for a single-compartment thickener is calculated using Figure D.3 (Perry et

al., 1997: 18-73). The installed cost include the raking mechanism (including drivehead and lift),

walkways and bridge of centerpier, cage, railings, and overflow launders.

Figure D.3: Installed cost for single-compartment thickeners

The calculations for the purchased equipment cost calculated from the equations and

correlations above are given in Table D.1.



Table D.1: Purchased equipment cost

Equipment Quantity Price per unit (R per unit) Total price

(R)

Leaching

Storage tanks for nitric acid 4 3,641,746.76 14,566,987.06

Slurry pumps 2 156,478.09 * 312,956.19

Thickener capacitance tank 1 15,462,463.06 15,462,463.06

Wash solution capacitance tank 1 4,339,077.16 4,339,077.16

CCD

Clarifier for train 2 1 13,708,772.19 13,708,772.19

Thickeners train 2 5 10,344,921.71 51,724,608.53

Slurry pumps 30 152,209.38 * 4,566,281.35

Ion exchange

Wash water capacitance tank 1 1,789,331.97 1,789,331.97

Eluant make-up tank 1 1,173,448.55 1,173,448.55

Regeneration make-up tank 1 1,173,448.55 1,173,448.55

Liquid pumps 14 126,342.70 * 1,768,797.80

Slurry pumps 6 99,958.79 * 599,752.72

Solvent extraction

Mixer-settlers 13 1,076,923.08 * 14,000,000.00

Liquid pumps 22 11,832.86 * 283,988.64

Organic make-up tank 1 358,988.91 358,988.91

Raffinate tank 1 956,138.27 956,138.27

OK-liquor storage tanks 1 956,138.27 956,138.27

Regen make-up 1 459,475.00 459,475.00

Organic storage tank 1 956,138.27 956,138.27

(NH4)2SO4 make-up tank 1 956,138.27 956,138.27

Caustic storage tank 1 956,138.27 956,138.27

Spent regen 1 956,138.27 956,138.27

Eluate capacitance tank 1 4,339,077.16 4,339,077.16

Demineralized water 1 956,138.27 956,138.27



Precipitation

Heat exchanger 1 12,736,721.48 12,736,721.48

Storage tank for solvent extraction recycle 1 769,550.32 769,550.32

Storage tank for stage 1 centrifuge recycle 1 263,056.52 263,056.52

Storage tank for stage 2 centrifuge recycle 1 88,392.18 88,392.18

ADU storage tank 1 1,457,967.02 1,457,967.02

Liquid pumps 6 24,649.08 * 147,894.46

Slurry pumps 14 21,907.19 * 306,700.62

Total purchased equipment cost (E) 153,090,705.31

* Average price per unit

The total capital investment is calculated using the total purchased equipment cost (E) and the

factors given in Table 5.1 for a solid-fluid processing plant. In Table D.2 the results of these

calculations are given.

Table D.2: Capital investment calculations

Solid-fluid processing plant factor

Calculated values (R million)

Direct costs

Purchased equipment cost (E) 153.09

Delivery, percent of purchased equipment 0.10 15.31

Subtotal: delivered equipment 168.40

Installation 0.39 65.68

Instrumental and controls 0.26 43.78

Piping 0.31 52.20

Electrical systems 0.10 16.84

Buildings 0.29 48.84

Yard improvements 0.12 20.21

Service facilities 0.55 92.62

Total direct cost 508.57



Indirect costs

Engineering and supervision 0.32 53.89

Legal expenses 0.04 6.74

Construction expenses 0.34 57.26

Contractors fee 0.19 32.00

Contingency 0.37 62.31

Total indirect cost 212.18

Total fixed capital 720.75

Working capital 0.75 126.30

Total capital invesment 847.05

D.2. Operating cost

The detailed operating cost calculations are given in Table D.3 with the following assumptions:

• 6 skilled operators per shift.

• 3 semi-skilled operators per shift.

• 5 general assistants per shift.

• The property is owned by AngloGold Ashanti, thus no rent is paid.

• Already existing market and buyer for the product ADU and gold.

• There is no research and development department on the plant.

• No royalties costs are included as no equipment or processes needs to be patented.



Table D.3: Detailed operating cost calculation

Production cost Suggested

factor

Rate or quantity per

year

Cost per rate or quantity

unit Unit

Calculated values (R

million per annum)

Variable production costs

Raw materials

Nitric acid 13844 1,500.00 R/ton 20.77

Sulphuric acid 143006 500.00 R/ton 71.50

Resin (Ambersep 400) 123 57,000.00 R/ton 7.03

Caustic soda (NaOH) 168 4,200.00 R/ton 0.71

Kerosene 283 5,000.00 R/m3 1.42

Isodecanol 9 22,410.00 R/ton 0.21

Alamine 9 57,260.00 R/ton 0.53

NH3 175 15,000.00 R/ton 2.62

Lime (CaO) 83017 139.00 R/ton 11.54

Na2CO3 84 900.00 R/ton 0.08

Subtotal: Raw materials 116.39

Operating labor

Skilled 48900 67.80 R/hr 3.32

Semi-skilled 24450 31.40 R/hr 0.77

General assistant 40750 20.20 R/hr 0.82

Call out feed 1630 33.60 R/hr 0.05

Subtotal: Operating labour 4.96

Operating supervision 0.15 of operating labour 0.74

Utilities

Water

Cooling 0 7.00 R/m3 0.00

Process 3586 7.00 R/m3 0.03

Electricity 24450000 0.40 c/kWh 9.78

Fuel 0 7.80 R/L 0.00

Refrigeration 0 148.32 R/GJ 0.00

Steam (175°C and 12 bar) 312442 50.00 R/ton 15.62



Waste water

Disposal 0 3.93 R/ton 0.00

Treatment 16820 3.93 R/ton 0.07

Waste treatment and disposal

Hazardous 100 1,075.33 R/ton 0.11

Non-hazardous 1000 266.98 R/ton 0.27

Subtotal: Utilities 25.87

Maintenance and repairs 0.07 of Fixed capital investment 50.45

Operating supplies 0.15 of Maintenance and repairs 7.57

Laboratory charges 0.15 of Operating labour 0.74

Royalties 0.04 of Total product cost without depreciation

Catalysts and solvents

(Magnafloc 90L) 215.975 22,500.00 R/ton 4.86

Total: Variable costs 211.59

Fixed charges (without depreciation)

Taxes (property) 0.02 of Fixed capital investment 14.42

Financing (interest) 0.105 of Fixed capital investment 75.68

Insurance 0.01 of Fixed capital investment 7.21

Rent 0 of Fixed capital investment 0.00

Subtotal: Fixed charges 97.30

Plant overhead costs 0.5

of operating labour, supervision and

maintenance 28.08

Total manufacturing costs 336.97

Administrative costs 0.2 of operating labour 0.99

Distribution and marketing

expenses 0 of total product cost 0.00

Research and development

costs 0 of every sales Rand 0.00

Total: Fixed cost 126.37

Total product cost 337.96



The total variable cost calculated in Table D.3 is used in the overall cash flow analysis but is

used as rand per ton ADU produced. Using the variable cost of R 211.59 million per annum and

a production rate of 1068.20 ton ADU per annum the variable cost per ton ADU produced is

calculated to be R 198 079.21. The final product cost accumulates to approximately

R 337.96 million per annum.

D.3. Revenue

Revenue for this operating plant is generated from the uranium and gold product sales. In

Table D.4 the price at which ADU is sold internationally and to NUFCOR as well as the trading

gold price is given. These spot prices were received on 26 October 2009 from William Manana

at South Uranium Plant.

Table D.4: Revenue calculations

Product International trading price Selling price to NUFCOR

ADU 45.5 $/lb 22.75 $/lb

Gold 254 552.77 R/kg N/A

D.4. Cash flow analysis

Using the calculations above a cash flow analysis is done and given in Table D.5. The cash

flow diagram in Figure 5.1 is achieved from the cash flow analysis in Table D.5.

Table D.5: Cash flow analysis.

Capital

Fixed capital R 720,751,040.60

Results Working capital R 126,299,831.88

ROI 60.49%

Costs

Fixed costs R 126,372,504.60 per annum

Payback (years) 3.98 Variable cost R 198,079.21 per ton

NPV R 2,363,686,680.54

Revenue from sales

ADU R 371,893.21 per ton

IRR 352%

Gold R 127,276,385.00 per ton Tax rate 28.0% per annum Discount rate 11.0% per annum Inflation rate 6.4% per annum



CCD construction (old production)

SX construction (no production)

Optimization (Half production)

Operation (full production)

Year 0 1 2 3 4 ADU production rate (ton/year) 624 534 1068 Gold production rate (ton/year) 1.44 1.08 2.16 Fixed capital R -720,751,040.60 Working capital R -126,299,831.88 Depreciation R -36,037,552.03 R -36,037,552.03 Inflation factor 1.00 1.06 1.13 1.20 Fixed costs R -84,248,336.40 R -143,065,806.97 R -152,222,018.62 Variable costs R -82,400,950.85 R -131,310,126.02 R -254,868,495.73 Revenue from sales for ADU R 232,061,360.12 R 224,866,045.87 R 478,514,945.60 Revenue from sales for gold R 183,277,994.40 R 155,616,213.26 R 331,151,301.82 Profit before tax R 248,690,067.27 R 106,106,326.14 R 402,575,733.07 Tax R -69,633,218.83 R -19,619,256.75 R -102,630,690.69 Profit after tax R 179,056,848.43 R 86,487,069.39 R 299,945,042.38 Cash flow R - R 179,056,848.43 R -847,050,872.48 R 86,487,069.39 R 299,945,042.38 Cumulative cash flow R - R 179,056,848.43 R -667,994,024.05 R -581,506,954.66 R -281,561,912.28 Discount factor 1.00 0.90 0.81 0.73 Discounted cash flow R - R 179,056,848.43 R -763,108,894.13 R 70,194,845.70 R 219,317,229.85 Cumulative discounted cash flow R - R 179,056,848.43 R -584,052,045.70 R -513,857,199.99 R -294,539,970.14 IRR Discount Factor 1.00 0.22 0.05 0.01 IRR Discounted Cash Flow R - R 179,056,848.43 R -187,547,866.97 R 4,239,908.18 R 3,255,739.08 IRR Cumulative Discounted Cash Flow R - R 179,056,848.43 R -8,491,018.53 R -4,251,110.36 R -995,371.27




Year 5 6 7 8 9 ADU production rate (ton/year) 1068 1068 1068 1068 1068 Gold production rate (ton/year) 2.16 2.16 2.16 2.16 2.16 Fixed capital Working capital Depreciation R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 Inflation factor 1.28 1.36 1.45 1.54 1.64 Fixed costs R -161,964,227.81 R -172,329,938.39 R -183,359,054.44 R -195,094,033.93 R -207,580,052.10 Variable costs R -271,180,079.46 R -302,437,433.47 R -307,001,883.24 R -326,650,003.76 R -364,301,054.04 Revenue from sales for ADU R 509,139,902.12 R 541,724,855.85 R 576,395,246.63 R 613,284,542.41 R 652,534,753.13 Revenue from sales for gold R 352,344,985.14 R 374,895,064.19 R 398,888,348.29 R 424,417,202.58 R 451,579,903.55 Profit before tax R 428,340,579.99 R 441,852,548.18 R 484,922,657.24 R 515,957,707.31 R 532,233,550.54 Tax R -109,844,847.83 R -113,628,198.92 R -125,687,829.46 R -134,377,643.48 R -138,934,879.58 Profit after tax R 318,495,732.16 R 328,224,349.26 R 359,234,827.78 R 381,580,063.83 R 393,298,670.95 Cash flow R 318,495,732.16 R 328,224,349.26 R 359,234,827.78 R 381,580,063.83 R 393,298,670.95 Cumulative cash flow R 36,933,819.88 R 365,158,169.14 R 724,392,996.92 R 1,105,973,060.75 R 1,499,271,731.70 Discount factor 0.66 0.59 0.53 0.48 0.43 Discounted cash flow R 209,803,003.91 R 194,785,175.97 R 192,061,608.68 R 183,791,247.17 R 170,662,714.29 Cumulative discounted cash flow R -84,736,966.24 R 110,048,209.73 R 302,109,818.41 R 485,901,065.58 R 656,563,779.87 IRR Discount Factor 0.00 0.00 0.00 0.00 0.00 IRR Discounted Cash Flow R 765,445.29 R 174,656.19 R 42,324.73 R 9,954.15 R 2,271.66 IRR Cumulative Discounted Cash Flow R -229,925.98 R -55,269.79 R -12,945.06 R -2,990.91 R -719.25




Year 10 11 12 13 14 ADU production rate (ton/year) 1068 1068 1068 1068 1068 Gold production rate (ton/year) 2.16 2.16 2.16 2.16 2.16 Fixed capital Working capital Depreciation R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 Inflation factor 1.75 1.86 1.98 2.11 2.24 Fixed costs R -220,865,175.43 R -235,000,546.66 R -250,040,581.65 R -266,043,178.87 R -283,069,942.32 Variable costs R -369,799,162.66 R -393,466,309.07 R -438,818,887.11 R -445,441,634.63 R -473,949,899.25 Revenue from sales for ADU R 694,296,977.33 R 738,731,983.88 R 786,010,830.85 R 836,315,524.02 R 889,839,717.56 Revenue from sales for gold R 480,481,017.38 R 511,231,802.49 R 543,950,637.85 R 578,763,478.67 R 615,804,341.30 Profit before tax R 584,113,656.61 R 621,496,930.63 R 641,101,999.94 R 703,594,189.18 R 748,624,217.29 Tax R -153,461,309.28 R -163,928,626.01 R -169,418,045.41 R -186,915,858.40 R -199,524,266.27 Profit after tax R 430,652,347.33 R 457,568,304.63 R 471,683,954.52 R 516,678,330.78 R 549,099,951.02 Cash flow R 430,652,347.33 R 457,568,304.63 R 471,683,954.52 R 516,678,330.78 R 549,099,951.02 Cumulative cash flow R 1,929,924,079.03 R 2,387,492,383.66 R 2,859,176,338.18 R 3,375,854,668.96 R 3,924,954,619.98 Discount factor 0.39 0.35 0.32 0.29 0.26 Discounted cash flow R 168,352,670.45 R 161,148,454.87 R 149,657,448.35 R 147,687,759.61 R 141,401,065.09 Cumulative discounted cash flow R 824,916,450.32 R 986,064,905.19 R 1,135,722,353.54 R 1,283,410,113.15 R 1,424,811,178.24 IRR Discount Factor 0.00 0.00 0.00 0.00 0.00 IRR Discounted Cash Flow R 550.74 R 129.56 R 29.57 R 7.17 R 1.69 IRR Cumulative Discounted Cash Flow R -168.50 R -38.94 R -9.37 R -2.19 R -0.51




Year 15 16 17 18 ADU production rate (ton/year) 1068 1068 1068 1068 Gold production rate (ton/year) 2.16 2.16 2.16 2.16 Fixed capital Working capital Depreciation R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 Inflation factor 2.38 2.54 2.70 2.87 Fixed costs R -301,186,418.63 R -320,462,349.42 R -340,971,939.79 R -362,794,143.93 Variable costs R -528,579,353.65 R -536,556,785.14 R -570,896,419.39 R -636,700,336.56 Revenue from sales for ADU R 946,789,459.48 R 1,007,383,984.89 R 1,071,856,559.92 R 1,140,455,379.76 Revenue from sales for gold R 655,215,819.15 R 697,149,631.57 R 741,767,207.99 R 789,240,309.31 Profit before tax R 772,239,506.35 R 847,514,481.90 R 901,755,408.74 R 930,201,208.57 Tax R -206,136,547.21 R -227,213,540.36 R -242,400,999.88 R -250,365,823.83 Profit after tax R 566,102,959.14 R 620,300,941.54 R 659,354,408.86 R 679,835,384.74 Cash flow R 566,102,959.14 R 620,300,941.54 R 659,354,408.86 R 679,835,384.74 Cumulative cash flow R 4,491,057,579.12 R 5,111,358,520.65 R 5,770,712,929.51 R 6,450,548,314.25 Discount factor 0.23 0.21 0.19 0.17 Discounted cash flow R 131,332,956.81 R 129,645,593.01 R 124,151,294.99 R 115,322,255.03 Cumulative discounted cash flow R 1,556,144,135.05 R 1,685,789,728.06 R 1,809,941,023.05 R 1,925,263,278.08 IRR Discount Factor 0.00 0.00 0.00 0.00 IRR Discounted Cash Flow R 0.39 R 0.09 R 0.02 R 0.01 IRR Cumulative Discounted Cash Flow R -0.12 R -0.03 R -0.01 R -0.00




Year 20 21 22 ADU production rate (ton/year) 1068 1068 1068

Gold production rate (ton/year) 2.16 2.16 2.16 Fixed capital Working capital R 126,299,831.88 Depreciation R -36,037,552.03 R -36,037,552.03 R -36,037,552.03 Inflation factor 3.25 3.46 3.68 Fixed costs R -410,717,799.17 R -437,003,738.32 R -464,971,977.57 Variable costs R -687,673,364.18 R -766,937,482.09 R -778,512,264.90 Revenue from sales for ADU R 1,291,104,973.60 R 1,373,735,691.91 R 1,461,654,776.19 Revenue from sales for gold R 893,495,797.20 R 950,679,528.22 R 1,011,523,018.03 Profit before tax R 1,086,209,607.45 R 1,120,473,999.73 R 1,229,693,551.76 Tax R -294,048,175.52 R -303,642,205.36 R -334,223,679.92 Profit after tax R 792,161,431.93 R 816,831,794.38 R 895,469,871.83 Cash flow R 792,161,431.93 R 816,831,794.38 R 1,021,769,703.71 Cumulative cash flow R 7,987,829,318.43 R 8,804,661,112.80 R 9,826,430,816.52 Discount factor 0.14 0.12 0.11 Discounted cash flow R 109,062,913.97 R 101,314,838.89 R 114,174,854.50 Cumulative discounted cash flow R 2,148,196,987.14 R 2,249,511,826.03 R 2,363,686,680.54 NVP

IRR Discount Factor 0.00 0.00 0.00 IRR Discounted Cash Flow R 0.00 R 0.00 R 0.00 IRR Cumulative Discounted Cash

Flow R -0.00 R -0.00 R 0.00



D.5. Economic sensitivity analysis

The results of the economic sensitivity analysis done is given in Tables D.6 and D.7.

Table D.6: Chance in variables

Revenue from a unit ADU (R/ton)

Inflation Fixed Cost (R/annum)

Variable Cost (R/ton ADU)

Tax Rate

20% 446,271.85 7.7% 151,647,005.52 237,695.05 33.6%

15% 427,677.19 7.4% 145,328,380.29 227,791.09 32.2%

10% 409,082.53 7.0% 139,009,755.06 217,887.13 30.8%

5% 390,487.87 6.7% 132,691,129.83 207,983.17 29.4%

0% 371,893.21 6.4% 126,372,504.60 198,079.21 28.0%

-5% 353,298.55 6.1% 120,053,879.37 188,175.25 26.6%

-10% 334,703.88 5.8% 113,735,254.14 178,271.29 25.2%

-15% 316,109.22 5.4% 107,416,628.91 168,367.33 23.8%

-20% 297,514.56 5.1% 101,098,003.68 158,463.37 22.4%

Table D.7: Economic sensitivity analysis in terms of percentage rise/fall of NPV

Revenue from a unit ADU Inflation Fixed Cost Variable Cost Tax Rate

20% 30.9% 16.6% -10.3% -16.2% -9.4%

15% 23.2% 12.2% -7.7% -12.2% -7.0%

10% 15.5% 8.0% -5.1% -8.1% -4.7%

5% 7.7% 3.9% -2.6% -4.1% -2.3%

0% 0.0% 0.0% 0.0% 0.0% 0.0%

-5% -7.7% -3.8% 2.6% 4.1% 2.3%

-10% -15.5% -7.4% 5.1% 8.1% 4.7%

-15% -23.2% -10.9% 7.7% 12.2% 7.0%

-20% -30.9% -14.3% 10.3% 16.2% 9.4%


299 Appendix E: Plant layout and positioning

Appendix E: Plant layout and positioning

Since this project is an upgrade of the existing South Uranium Plant of AngloGold Ashanti

located near Orkney in the North-West Province, the old plant layout should be studied to

analyze available space to build new processing equipment. Figure E.1 is an air photo taken

from the program Google Earth® and display the old layout of the plant.

Figure E.1: Google Earth air photo

Figure E.2 and E.3 is the plant layout received from William Manana at AngloGold Ashanti for

the existing plant.

CCD

Leaching

Solvent extraction

IX and precipitation

Sulphuric acid storage

South Uranium Plant (SUP)



Figure E.2: Existing plant layout of South Uranium Plant with number legend

LEGEND

1. Water treatment salt 2. Water treatment salt 3. Floc N300 750kg 25kg 4. Soda Ash

Copper Sulphate Floc N300 25kg

5. Lime 6. Manganese 7. Manganese 8. Coal 9. Diesel 10. Ammonia 11. Caustic 12. Pegasol, Armeen 380, Isodec 13. Steel Balls 14. Armeen drums 15. Resin 16. Resin 17. Manganese



Figure E.3: Existing plant layout of South Uranium Plant with colour legend


302 Appendix F: MSDS information

Appendix F: MSDS information

The MSDS forms for the following substances are provided on the compact disc (CD)

provided with the report:

• Alamine® 336 (Cognis corporation, 2007).

• Ammonia gas (BOC Gases, 1996).

• Ammonium diuranate (ADU) (International Bio-Analytical Industries, Inc, 2006).

• Caustic soda (ScienceLab, 2009).

• Ferrous sulphate (ScienceLab, 2009).

• Isodecanol (BASF, 2006).

• Kerosene (ScienceLab, 2009).

• Magnafloc 90L (ACAT, 2006).

• Nitric acid (ScienceLab, 2009).

• Nitric oxide (Air products, 1998).

• Potable water (ScienceLab, 2009).

• Silicon dioxide (ScienceLab, 2009).

• Sulphuric acid (ScienceLab, 2009).

• Uranium (British Geological Survey, 2007).

ipl engineering design report for cemi427

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