optimizing membrane distillation process for triethylene...

i

Optimizing Membrane Distillation Process for Triethylene Glycol

Separation from Gas Separation Plant Waste Stream

by

Pham Minh Duyen

A thesis submitted in partial fulfillment of the requirements for the

degree of Master of Engineering in

Environmental Engineering and Management

Examination Committee: Prof. Chettiyappan Visvanathan (Chairperson)

Dr. Thammarat Koottatep

Dr. Romchat Rattanaoudom (External Expert)

Nationality: Vietnamese

Previous Degree: Bachelor of Engineering in Environmental Engineering

Ho Chi Minh City University of Technology

Vietnam

Scholarship Donor: Greater Mekong Subregion (GMS) Scholarship

Asian Institute of Technology

School of Environment, Resources and Development

Thailand

May 2015

ii

Acknowledgments

My master thesis would not have been possible to achieve without the support of many

kind people. It is my pleasure to acknowledge everyone who has supported me to achieve

this thesis and also my master degree.

Above all, I would like to express my heartfelt gratitude to my advisor, Prof. C.

Visvanathan. This thesis would never have been achieved without his encouragement and

kind patience along this challenging journey. I am truly thankful for his helpful advice and

technical suggestions. By having Prof. C. Visvanathan as advisor, I have improved the

confidence in both academic and personal sides.

I also would like to show my appreciation to my thesis’s committee members, Dr.

Thammarat Koottatep and Dr. Romchat Rattannaoudom. Their useful advice, support and

encouragement have been invaluable upon the completion of my thesis.

I am sincerely thankful to EEM faculties and staffs. Mr. Panupong and Mr. Nimitra are two

technicians of EEM who I am most grateful to. The experimental set-up would not be

controlled, fixed and improved without their technical support. I wish to deliver my grateful

appreciation to Mr. Chaiyaporn who provided laboratory facilities and analysis information

together with many important advice. Furthermore, I want to thank Ms. Suchitra, Ms.

Chanya, Ms. Salaya amd Ms. Orathai for their kind support and encouragement. Especially,

I am grateful to Ms. Suchitra who always took care of me in both academic and personal

life, and encouraged me to move on.

My special thanks also go to Prof. C. Visvanathan’s research group. I am happy to have

worked alongside this highly motivated team during my master thesis. Mr. Thusitha and

Mr. Jacob acted as helpful mentors with many critical technical input for my study.

For my colleagues, Kung, My, Thi, Aashik, Ju Ju, Milk and Mr. Park, thank you all for

your help and support. It is very appreciated to your sharing important information and

suggestion, especially our great friendship.

I thankfully acknowledge and appreciate the Royal Thai Government for the financial

support with Loom Nam Khong Pijai (GMSARN) scholarship in order for me to pursue

the master program. My sincere thanks are expressed to PTT Public Company Limited

(Thailand) for providing wastewater and conducting wastewater analysis throughout this

research. Especially, my appreciation is expressed for the AIT-PTT-II project for the

research grant.

Last but not least, to my dear family, all my expression of gratitude does not suffice. I am

grateful for their love, encouragement and support. I want to say thank you for their

understanding and being my home no matter what I did, especially during my away-from-

home mission.

iii

Abstract

In natural gas processing, triethylene glycol (TEG) is used as a dehumidifying agent to

absorb and remove water content in the process called as dehydration. TEG should be

recovered from wastewater and could be reused. Hydrophobic membrane distillation (MD)

has a great potential to concentrate TEG in wastewater. This study mainly focused on

optimizing the operational condition of membrane distillation process for concentrating

TEG from wastewater. Two scale of hollow fiber MD were investigated (0.25 m2 and 2

m2). Each membrane module was tested with three solutions: pure water, synthetic TEG

and real wastewater alternately.Energy consumption analysis in all experiments was

evaluated in term of the ratio of energy consumption/permeate flow. Fouling analysis was

conducted to evaluate the quality of cleaning process.

For bench scale study, the optimum condition in terms of energy consumption/Qp and

permeate flux were achieved at feed flow rate of 2.4 (L/min), feed temperature of 70oC,

and sweeping gas velocity of 4.7 m/s (gas inlet flow rate of 0.255 L/min.fiber). At this

condition, pure water flux achieved was 3.14 kg/m2.h, the ratio of energy consumption/Qp

was 1.09 kWh/kg. During concentrating synthetic TEG from 10 to 45 %, the permeate flux

was in the range from 2.1 to 2.61 kg/m2.h with the ratio of energy consumption/Qp was

approximately 1.29 kW/kg. The permeate flux of real wastewater investigation when

concentrating TEG from 9.69 to 50% were in the range from 2.4 to 1.6 kg/m2.h with the

ratio of energy consumption/Qp was 1.4 kWh/kg. It could concentrate TEG in real

wastewater till 98.01%. In the total resistance, membrane resistance, boundary layer

resistance, fouling resistance contributed 69.2 %, 7.6%, 23.2 % respectively. The

irreversible resistance accounted for 2%.

The optimum condition in pilot scale study was decided at the gas velocity of 8.07 m/s (gas

inlet flow rate of 0.44 L/min.fiber), feed inlet temperature of 70oC. Feed flow rate and

sweeping gas temperature have less effect. The flux achieved was 1.94 kg/m2.h. At this

condition, the ratio of energy consumption/Qp was 0.51 kW/kg. The permeate flux was

stable around 1.8 kg/m2.h when the synthetic TEG concentration was less than 30 %. When

the synthetic TEG concentration was in the range from 30 to 50%, permeate flux was about

1.6 kg/m2.h (average). At higher TEG concentration (>50%), the flux reduced to less than

1.2 kg/m2.h. When TEG concentration was in the range from 10 to 60%, the ratio of energy

consumption/Qp was less than 0.7 kW/kg. Pilot scale MD module has ability to concentrate

synthetic TEG till 98.7%. Investigating pilot scale MD module with real TEG wastewater,

the permeate flux was higher than 1.7 kg/m2.h with the TEG concentration less than 15 %.

The flux reduced to the range between 1.4 and 1.65 kg/m2.h at higher TEG concentration

from the hour of 5th and 28th (TEG concentration less than 50%). It reached 0.9 kg/m2.h at

the hour of 40th (final TEG concentration was about 69.4 %). Pilot scale SGMD module

had ability to concentrate TEG in real wastewater until 99.1%. The contributions of

membrane resistance, boundary layer resistance and fouling resistance into the total

resistance were 60.2 %, 10.4 % and 29.4% respectively. After cleaning with chemical

agents, irreversible fouling resistance remained 1.3% total resistance. To concentrate real

wastewater from 10 to 45 % TEG concentration, the required energy ratio was 0.62 kW/kg.

The energy ratio in real wastewater and synthetic TEG were similar at 0.95 kW/kg when

concentrating TEG continuously up to 65%.

iv

Graphical Abstract

1

Table of Contents

Chapter Title Page

Title Page

Acknowledgements

Abstract

Graphical Abstract

Table of Contents

List of Tables

List of Figures

List of Abbreviations

i

ii

iii

iv

v

ix

xi

xiv

1

2

3

4

Introduction

1.1 Background

1.2 Objectives of the Study

1.3 Scope of the Study

Literature Review

2.1 Natural Gas Industry Overview

2.2 Triethylene Glycol

2.3 Membrane Science and Technology

2.4 Membrane Distillation

2.5 Material of Membrane and Module Fabrications Using for

Distillation Process

2.6. Operational Processes of Membrane Distillation

2.7 Temperature Polarization

2.8 Concentration Polarization

2.9 Membrane Fouling

2.10 Membrane Cleaning

2.11 Operating Variables Affecting MD Process

2.12 Effects of Membrane Parameters on MD Process

2.13Advantages and Limitations of Membrane Distillation

Technology

2.14 Application of Membrane Distillation

2.15 Research Gap

Methodology

3.1 Methodology Overview

3.2 Experimental Materials

3.3 Experimental Methods

3.4 Experimental Analysis

Results and Discussions

4.1 TEG Wastewater Characterization

4.2 Performance of Pre-treatment Unit

4.3 Optimizing Operation of Bench Scale Hollow Fiber SGMD

4.4 optimum condition of pilot scale SGMD and evaluate energy

4.5 Full Scale SGMD Plant Design

1

2

2

2

3

8

12

13

17

18

25

26

27

29

29

33

35

36

37

39

39

41

44

51

54

54

55

56

79

104

2

5

Conclusions and Recommendations

5.1 Conclusions

5.2 Recommendations for Future Study

Reference

Appendix A

Appendix B

Appendix C

Appendix D

Appendix E

Appendix F

108

108

110

112

118

123

151

163

174

200

3

List of Tables

Table

Title Page

2.1 Components of Natural Gas 3

2.2 Properties of Natural Gas 4

2.3 Properties of Triethylene Glycol 9

2.4 TEG Applications 10

2.5 General Characteristics of Membrane Process 16

2.6 Effect of Variables on Permeate Flux in MD Process 32

2.7 Effect of Membrane Parameter on Permeate Flux in MD Process 32

2.8 Typical Fields of MD Application 37

2.9 Overview of Various TEG Separation Processes 37

3.1 MF and UF Membrane Properties in Pre-treatment Unit 41

3.2 Hollow Fiber Membrane Distillation Specification 42

3.3 Membrane Distillation Module Specification 43

3.4 Cleaning Chemicals Used in this Study 44

3.5 Membrane Cleaning Procedure 51

3.6 Analytical Parameters and Methods 53

4.1 Wastewater Analytical Results 54

4.2 Analytical Results of Pre-treating Samples 55

4.3 The Ratio of Energy Consumption/Qp at Different Scenarios of Pure

Water Test

61

4.4 PWF Comparison with Other Studies on SGMD 63

4.5 Membrane Surface Temperature and Temperature Polarization

Coefficient

64

4.6

Experimental Membrane Distillation Coefficient and Membrane

Resistance

64

4.7 TEG Concentration in Synthetic TEG Test 65

4.8 Membrane Resistance and Boundary Layer Resistance 68

4.9 Total Membrane Boundary Layer and Fouling Resistance Calculated

from Fouled Permeate Flux (batch operation)

72


from Fouled Permeate Flux (continuously-fed operation)

73

4.11 Comparison of Fouling and Other Resistance between Continuously-fed

and Batch Operation

74

4.12 Comparison of Fouling and Other Resistance between SGMD and

DCMD Bench Scale Hollow Fiber Membrane Distillation

(Continuously-fed)

75

4.13 Experimental Results of PWF Investigation on Pilot Scale Module 81

4.14 Membrane Resistance and Membrane Coefficient 81

4.15 Membrane Resistance and Boundary Layer Resistance 89


from Fouled Permeate Flux

92


from Fouled Permeate Flux (continuously-fed operation)

94

4.18 Comparison of Fouling and Other Resistance between Continuously-fed

and Batch Operation

95

4

4.19 Comparison of Fouling and Other Resistance between Normal

Condition and 0ptimum Condition of SGMD Hollow Fiber Membrane

Distillation (Continuously-fed)

96

4.20 Summary of Financial Analysis for Overall System with ±30% Variation 107

5

List of Figures

Figure

Title Page

2.1 Natural gas use by sector 5

2.2 Products of gas separation plant and applications 6

2.3 Natural gas processing 7

2.4 Triethylene glycol wastewater steam of PTT’s gas separation plant 8

2.5 Break down of TEG application in US 9

2.6 Vapor-liquid interface in membrane distillation 14

2.7 Four configurations of membrane distillation process 14

2.8 Dusty gas model 21

2.9 Poiseuille type of flow inside a pore of SGMD process 23

2.10 Temperature and concentration profile in MD process 26

2.11 Concentration and temperature polarization in MD process 27

2.12 Profile of temperature in fouling case 28

3.1 Experimental study plan 40

3.2 Experimental materials using in this research 41

3.3 Schematic diagram of pre-treatment unit 45

3.4 Image of pre-treatment system 45

3.5 Cross flow mode description in sweep gas membrane distillation 46

3.6 Bench scale hollow fiber sweep gas membrane distillation 46

3.7 Image of bench scale hollow fiber SGMD system 46

3.8 Pilot scale hollow fiber sweep gas membrane distillation unit 47

3.9 The principle of sweeping gas membrane distillation configuration 48

3.10 Determination process of fouling resistances 52

4.1 Removal efficiency of pre-treatment system 56

4.2 Rejection result of hollow fiber SGMD at gas velocity of 3.1 m/s 57

4.3 Rejection result of hollow fiber SGMD at gas velocity of 5.3 m/s 58

4.4 Permeate flux at different feed inlet flow rate and sweeping gas velocity at

feed temperature of 70oC

59

4.5 Permeate flux at different feed inlet temperature and sweeping gas velocity at

feed flow rate of 2.4 L/min

59

4.6 Energy consumption/Qp at feed flow rate of 2.4 L/min 61

4.7 Permeate flux and TEG concentration at synthetic TEG 10% initial

concentration

65


concentration

66


concentration

67

4.10 Experimental result of continuously-fed synthetic TEG investigation 67

4.11 Increasing of boundary layer resistance 69

4.12 Experimental results of batch experiment with real wastewater. 70

4.13 Experimental result of continuously-fed real wastewater investigation 71

4.14 Permeate flux and TEG concentration during 210 hrs feeding continuously 71

4.15 Classification of types of resistances in SGMD batch operation 73

4.16 Resistance classification in SGMD continuously-fed real wastewater 74

6

4.17 Energy Consumption/Qp during concentrating TEG at 10% initial

concentration

76


concentration

76


concentration

77

4.20 Energy Consumption/Qp during concentrating TEG for 40 hours 77

4.21 Energy Consumption/Qp of batch experiment with real wastewater. 78

4.22 Energy Consumption/Qp of continuously-fed experiments with real

wastewater.

79

4.23 Pure water flux at different sweeping gas inlet flow rate of pilot scale study 80

4.24 Rejection results for pilot scale SGMD process of pilot scale study 82

4.25 Permeate flux and TEG concentration synthetic TEG 10% initial

concentration of pilot scale study

83



83



84

4.28 Permeate flux and TEG concentration synthetic TEG 60% initial concentration

of pilot scale study

84



85



85



86



87



87



88

4.35 Experimental result of continuously-fed synthetic TEG investigation of pilot

scale study

89

4.36 Proportion of Boundary Layer Resistance and Membrane Resistance of pilot

scale study

90

4.37 Experimental results of batch experiment with real wastewater of pilot scale

study

91

4.38 Classification of types of resistances in SGMD batch operation of pilot scale

study

92

4.39 Experimental result of continuously-fed real wastewater investigation of pilot

scale study (40 hours)

93

4.40 Experimental result of continuously-fed real wastewater investigation of pilot

scale study (72 hours)

94

4.41 Classification of resistances in SGMD continuously-fed real wastewater of

pilot scale study

95

4.42 Inlet velocity distribution in hollow fiber module 97

4.43 Local velocities inside membrane module 98

7



99



99



100



100



101

4.49 Energy Consumption/Qp during concentrating TEG for 40 hours of pilot scale

study

102

4.50 Energy Consumption/Qp of batch experiment with real wastewater of pilot

scale study

102

4.51 Energy Consumption/Qp of continuously-fed experiments with real

wastewater of pilot scale study

103

4.52 Membrane distillation modules arrangement and pipeline levels 105

4.53 Sketch diagram of full scale SGMD plant 106

8

List of Abbreviations

AGMD Air gap membrane distillation

AIT Asian Institute of Technology

BOD Biochemical oxygen demand

Bw Membrane distillation coefficient

COD Chemical oxygen demand

CPC Concentration polarization coefficient

D Diffusion coefficient

DCMD Direct contact membrane distillation

dp Membrane pore size

E Energy efficiency

EE Evaporation efficiency

EG Ethylene glycol

FS Flat sheet

h Heat transfer coefficient

HF Hollow fiber

Jw Permeate flux

kb Boltzmann constant

Kn Knudsen number

LEP Liquid entry pressure

LPG Liquefied petroleum gas

MD Membrane distillation

MF Microfiltration

NF Nanofiltration

NGL Natural gasoline

P Total pressure

Pa Air pressure

PE Polyethylene

Pm Mean pressure within membrane pore

PP Polypropylene

PTFE Polytetrafluoroethylene

PTT PTT Public Company Limited PVC Polyvinylchloride

pw Vapor pressure

PWF Pure water flux

Q Heat flux

Qp Permeate flow rate

r Membrane pore radius

RO Reverse osmosis

Rw Membrane distillation resistance

SGMD Sweeping gas membrane distillation

T Absolute temperature

TDS Total dissolved solids

TEG Triethylene glycol

TPC Temperature polarization coefficient

TSS Total suspended solids

U Overall heat transfer coefficient

UF Ultrafiltration

9

VMD Vacuum membrane distillation

ΔHv Letent heat for evaporation

λ Mean free path

σw Collision diameter of water molecule

1

Chapter 1

Introduction

1.1 Background

Natural gas is considered as a very important non-renewable energy source. Biogenic and

thermogenic are two main mechanisms to generate natural gas over a long time. Natural

gas has widely applications which are mainly based on heat energy that is generated from

burning process. Those applications can be divided in four intensive sectors of society:

transportation, domestic use (heating and cooking), power generation (electricity) and

industrial production (i.e. fertilizer). As a type of fossil fuel, the crude natural gas is not a

pure source. In natural gas, besides the main component is methane gas (CH4), there are

plenty of other components and impurities as other species of hydrocarbons (alkane),

hydrogen sulfide (H2S), carbon dioxide (CO2), nitrogen (N2), moisture content. High

percentage of water vapor in natural gas can result in freezing pipelines, reduces the fuel’s

calorific value or other problems in application process.

Glycol is a homologous series of di-hydroxyl alcohols which obtains: ethylene glycol

(MEG), diethylene glycol (DEG), triethylene glycol (TEG) and tetraethylene glycol

(TREG). TEG is a co-product of MEG production process. It is a colorless and odorless

chemical which is low-volatility, water solubility, high viscosity and high boiling point.

On health risk aspect, TEG does not cause cancer. The path ways of human exposure are

inhalation (minimal risk due to low volatile property), dermal (skin or eyes irritation), oral

(adverse effect at lethal amount). On environmental risk aspect, TEG is a nontoxic

compound to aquatic life. In water and soil, the concentration of TEG is very low due to

the biodegradable nature (Dow, 2014).

In natural gas processing, triethylene glycol (TEG) is used as a dehumidifying agent to

absorb and remove water content in the process which is called as dehydration. There are

two source of TEG wastewater generated from dehydration unit. The first source is

generated from the condenser of TEG recovery system. It has a TEG concentration of 0.1

% by volume with the total volume generated of 19 m3 per day (PTT-GSP., 2012). The

second source comes from TEG trap of natural gas after crossing dehydration unit, the

concentration of TEG in this wastewater is various from 5-20%. The present of BTEX in

TEG wastewater resulted in contribution of air pollution and is considered as carcinogenic

source. Thus, TEG wastewater is a hazardous waste. Moreover, there are some other

pollutants in TEG wastewater such as suspended solids (SS), total dissolved solid (TDS),

oil, grease and heavy metal.

The first type of TEG wastewater (containing low TEG concentration) can be treated by

the conventional wastewater treatment process. In literature, nearly 98% TEG was removed

from this process (Alberta-Environment, 2010). The second source, containing very high

TEG concentration, is currently incinerated by a licensed company (PTT-GSP., 2012).

During the dehydration process, the physical properties of TEG do not change. Thus, TEG

should be recovered from wastewater and could be reused.

Membrane distillation (MD) technology is thermally-driven process (Khayet and

Matsuura, 2011) which has been developed more than 50 years ago. The difference in vapor

2

pressure between both sides of membrane is the driving-force of MD process. MD process

takes place once the partial vapor pressure of volatile compound of feed side is higher than

that in permeate side. In this process, hydrophobic membrane responses as a barrier that

only allows vapor to cross the pores. The separation of liquid - vapor is happened at the

entrance of each pore. Membrane distillation process has some significant advantages such

as less energy consumption, high selectivity, nearly 100 % rejection of non-volatile

material and less fouling condition.

At the pressure of 760 mmHg, the boiling point of TEG is 288oC (Dow, 2014) while water

boils at 100oC. In a mixture of liquid, the higher boiling point temperature leads to lower

partial vapor pressure. Consequently, the vapor pressure of TEG is always lower than that

of water. Thus, it is suitable to use MD for concentrating and recovering TEG from

wastewater stream.

1.2 Objectives of the Study

The objective of this study was to recover and concentrate TEG from wastewater. Then, it

could be reused in the process. To achieve this objective, three following objectives were

proposed and were achieved.

1. Optimizing operation of bench scale sweep gas membrane distillation system to

separate TEG from synthetic and real wastewater.

2. Scaling up the optimum condition to pilot scale for evaluating energy consumption

of the process.

3. Designing of full scale membrane distillation plant using all studied parameters.

1.3 Scope of the Study

The potential of low-cost technology on concentrating TEG from wastewater using

membrane distillation process (SGMD configuration) was studied in this thesis. Bench

scale hollow fiber SGMD unit (membrane surface area of 0.255 m2) was investigated in

the first phase. The optimum condition was selected based on the performance and energy

consumption of MD process. The results from bench scale study were applied in pilot scale

SGMD unit (membrane surface area of 2 m2) in the second phase to scale up the optimum

condition of this unit. Base on the experimental study, a full scale SGMD plan was designed

to treat real TEG wastewater from gas separation plant. This full scale MD plan had the

treatment capacity of 1 m3/day. Hence, the scopes of this study were as following:

1. Bench scale (0.255 m2) and pilot scale (2 m2) hollow fiber SGMD studies were

conducted.

2. Permeate flux and energy consumption are two factors that were used to evaluate

the performance of both bench scale and pilot scale unit. The variables include: feed

flow rate, feed concentration, feed temperature and sweep gas inlet flow rate.

3. Both synthetic TEG wastewater and real TEG wastewater were used in this study.

3

Chapter 2

Literature Review

2.1 Natural Gas Industry Overview

Natural gas is now considered as a vital fossil fuel in human’s life to generate non-

renewable energy. There are two main mechanisms to generate natural gas: biogenic and

thermogenic. In biogenic mechanism, methanogenic microorganisms in marshes were

responded. Thermogenic gases were produced by buried organic compounds in the deep

layer underground under intense heat and pressure over million years of time.

2.1.1 Natural gas properties

Methane is the primary hydrocarbon compound in natural gas. Commonly, natural gas also

consists of other alkanes (paraffinic hydrocarbon), and non-hydrocarbons such as carbon

dioxide, nitrogen, hydrogen sulfide and water content. The percentage of each compound

is shown in Table 2.1.

Table 2.1 Components of Natural Gas (Ibrahim, 2010)

Components

(IUPAC name) Molecular formula Percentage (%)

Methane CH4 >85

Ethane C2H6 3-8

Propane C3H8 1-2

Butane C4H10 <1

Pentane C5H12 <1

Carbon Dioxide CO2 1-2

Nitrogen N2 1-5

Hydrogen Sulphide H2S <1

Helium He <0.5

2.1.1.1 Chemical and physical properties

Natural gas is lighter than air. It is odorless, colorless and tasteless. The Wobbe index is

used as an indicator to evaluate the gas quality before selling. This is a ratio of calorific

value to the specific gravity. It is measured by the consideration of the heat input on a

typical appliance a given gas pressure.

2.1.1.2 Specific gravity

The specific gravity of natural gas is measured by Equation 3.1 below

http://en.wikipedia.org/wiki/International_Union_of_Pure_and_Applied_Chemistry_nomenclature

http://en.wikipedia.org/wiki/Molecular_formula

4

γg =M

Mair (2.1)

Where 𝛾𝑔 is the specific gravity of natural gas, 𝑀𝑎𝑖𝑟 is the molecular weight of air, 𝑀 is

the molecular weight of the mixture of natural gas.

Table 2.2 Properties of Natural Gas (Mokhatab et al., 2006)

Properties Unit Value

Relative molar mass 17-20

Carbon content % 73.3

Hydrogen content % 23.9

Oxygen content % 0.4

Hydrogen/carbon atomic ratio 3.0-4.0

Relative density at 15oC 0.72-0.81

Boiling point oC -162

Auto ignition temperature oC 540-560

Octane number 120-130

Methane number 69-99

Stoichiometric air/fuel ratio 17.2

Methane concentration % 80-99

2.1.1.3 Ideal and real gas laws

The volume of ideal gas is always higher than the volume of real gas due to super

compressible nature. The gas deviation factor (Z) is the ratio between real gas volume and

ideal gas volume at the given pressure and temperature. The real gas volume is calculated

by using Equation 2.2.

PV = ZnRT (2.2)

Where P, V, Z, n, R, T are pressure, volume, compressibility, number of kilo-moles of the

gas, gas constant and absolute temperature respectively. At low pressure and high

temperature (close to ideal condition), the value of Z is close to 1.

2.1.1.4 Gas formation volume factor

The gas formation volume factor (Bg) is measured by the ratio of the volume of 1 mole of

gas at a given condition (pressure and temperature) to the volume of 1 mole of gas at the

standard condition.

Bg = 0.3507ZT

P (2.3)

5

Where Bg is the gas formation factor (m3/Sm3), Z is the compressibility factor, P is pressure

(kPa), T is temperature (oK).

2.1.1.5 Gas density

The ratio of mass per volume of gas is the definition of gas density (𝜌𝑔). It is calculated

based on gas law.

ρg = 1.224γg

Bg (2.4)

Where 𝜌𝑔, Bgare in kg/m3 and m3/Sm3 respectively

2.1.1.6 Gas viscosity

Due to higher compressibility of natural gas in comparison with that of oil, water or rock,

the viscosity of gas is very low. It makes gas become easier to store in tank or reservoir.

2.1.2 Applications of natural gas

Most of applications of natural gas are based on heat energy generation. It can be used in

various sectors: electric power generation, hydrogen production, transportation, domestic

use, industrial use (i.e. fertilizer, steel, plastic).

Figure 2.1 Natural gas use by sector (Ibrahim, 2010)

Because natural gas is a mixture of various gases, each product from gas separation plant

is used with appropriate application. Figure 2.2 show the applications of natural gas’s

components.

Electric power

30%

Industrial

34%

Residental

20%

Commercial

13%

Others

3%

http://naturalgas.org/overview/uses-electrical/

6

2.1.3 Contaminants in raw natural gas

As mentioned earlier, natural gas is a mixture of various gases (hydrocarbons and non-

hydrocarbons). Some of them are considered as impurities which could be divided in three

main parts (Ibrahim, 2010).

2.1.3.1 Water vapor

This is a common impurity presenting in natural gas, and easy to be removed by

dehydration unit. The mixture of water vapor and H2S has resulted in corrosion of pipeline

system. Moreover, water and heavy hydrocarbons also produce solid hydrates which

causing pipeline clogging.

Figure 2.2 Products of Gas Separation Plant and Applications

2.1.3.2 Acidic gages

There usually are H2S and CO2 and SO3. H2S is toxic if burned. If the water presents, it

will combine with water to cause of corrosion of pipeline. CO2 leads to lower heating value.

2.1.3.3 Heavy hydrocarbons

It is undesirable to present in natural gas. Applicable equipment are not designed for this

type of hydrocarbons. In gas pipeline, it is difficult to have two phases flow: liquid and gas.

Raw Natural

Gas

Natural Gas

Processing

C2H6

C3H8

Petrochemical industry phase 1

NGL

Oil products and fuel for industries

Petrochemical industry phase 2

Raw material for food preservation CO2

CH4

Fuel for electric power generation

Fuel for industries

Natural gas for transportation

Raw material for fertilizer

LPG Fuel for cooking and vehicles

NGL: Natural Gas Liquid

LPG: Liquefied Petroleum Gas

7

2.1.4 Natural gas processing

Natural gas processing is an industrial process which uses raw natural gas as influent

material. There are two objectives of natural gas processing. The first is to remove

contaminants such as water vapor, excess hydrocarbon liquid, acidic gages to achieve the

marketable natural gas. The second purpose is to control the delivery pressure in the

distribution pipeline. The schematic of natural gas processing is shown in Figure 2.3.

Figure 2.3 Natural Gas Processing

2.1.5 Wastewater from gas separation plants

One of the most important units in natural gas processing is dehydration unit. In this step,

water vapor in natural gas stream is removed until its pressure reaches the value under the

dew point of natural gas in pipeline. The dew point of a gas is understood as a typical value

of temperature at which water vapor in this gas is condensed.

There are three purposes of dehydration unit: to avoid hydrate formation (water steam at

low temperature and high pressure in pipeline), to prevent corrosion of pipeline, and to

ensure that the water content in commercial gas will not create any problem for all

applications at downstream.

Inlet Gas Compression

Dehydration

Recovery of Natural Gas

Liquid

Sales Gas Compression

Transported Raw Gas

Phase Separation

Gas Treating

Water

Solids

Condensate to Stabilization Unit

Acid Gas to Sulphur Recovery Unit

Water

NGL to Fractionation

Distribution pipeline

8

In dehydration unit, Triethylene glycol (TEG) is usually used as a solvent in absorption

process. TEG in liquid form contacts directly with the wet gas. Water vapor transfers from

gas phase to liquid phase in this process (at high pressure and low temperature).

As presented in Figure 2.4, there are two sources of TEG wastewater from dehydration

unit. The first source is from TEG regeneration unit. This wastewater contains TEG of 0.1

% with the flow rate of 10 m3 /d. It will be sent to wastewater treatment plant to treat as an

industrial wastewater.

TEG

Cooler

Gas Dehydration Unit

TEG Filter TEG

Carbon Filter

TEG

After Filter

TEG Re-boiler

TEG

Still Column

Condenser

TEG Surge Tank

Contractor

Overhead Filter

Separator

0.01% Water

Gas Dehydration Unit

TEG Trap

TEG Trap0.01% Water

Dehydrated Gas

0.13% Water

0.13% Water

Feed Gas

Train 2

Feed Gas

Train 1

Gas

Gas Pipe Line

Wastewater 2

1 CMD

(TEG ~8-20%)

Wastewater 1

10 CMD

(TEG ~0.1%)

TEG Train 1

TEG Recovery Unit

Figure 2.4 Triethylene glycol wastewater steam of PTT’s gas separation plant

The second source is from TEG trap of natural gas after passing to dehydration unit. It has

the TEG concentration of 5-20 % and flow rate of 1 m3/d (PTT-GSP., 2012). This type of

wastewater, PTT is currently sending to Better World Green Company to treat with the

treatment cost of 4,500 Baht/m3. In this wastewater, besides TEG, it also contains BTEX

which are considered as carcinogenic compounds and contribute to air pollution.

2.2 Triethylene Glycol

2.2.1 Properties and applications

TEG formula: C6H14O4

Molecular formula: HO-CH2-CH2O-CH2-CH2O-CH2-CH2-OH

Triethylene glycol (TEG) is a co-product of the two reactions. The first is the oxidation of

ethylene which takes place at high temperature with the presence of silver oxide catalyst.

The second is hydration reaction of ethylene oxide to triethylene glycol.

9

TEG is a transparent chemical which is colorless, low volatility, high viscosity and water

soluble. TEG is odorless at normal conditions and sweet at high vapor concentration. Its

properties are familiar with other glycols in hydroxyl group. However, it is preferential to

use TEG in some applications which require higher boiling point, higher viscosity, higher

molecular weight and lower volatility than other glycols. In all applications of TEG, the

solubility is the most important characteristics.

Table 2.3 Properties of Triethylene Glycol (Dow, 2014)

Property Units Value

Auto-ignition Temperature oC 349

Boiling Point at 760 mmHg oC 288

Freezing Point oC -4.3

Heat of Vaporization kJ/gmol 62.5

Molecular Weight g/mol 150.17

Specific Gravity at 20 oC - 1.1255

Viscosity at 20 oC mPs 49

Surface Tension mN/m 45.5

Vapor pressure at 20 oC kPa <0.001

Solubility of Water in Triethylene Glycol at 20°C wt% 100

Solubility in Water at 20°C wt% 100

Figure 2.5 Break down of TEG application in US (Dow, 2014)

Gas dehydration

54%

Solvent

10%

Plasticizer

11%

Polyurethanes

9%

Humectant

4%

Others

12%

10

Table 2.4 TEG Applications (Dow, 2014)

Based on TEG Property Application

Hygroscopicity Dehydration of natural gas

Moisturizing and adhesives

Plasticizer Safety glass

Separation membranes (i.e. silicone rubber)

Ceramic materials

Low Volatility Gas dehydration

Solvent Steam-set printing inks

Aromatic and paraffinic hydrocarbons separation

Cleaning compounds

Cyanoacrylate and polyacrylonitrile

Resin impregnates

Chemical Intermediate Unsaturated polyester resin

Thermoplastic polyurethanes

Silicone compounds

Emulsifiers

Lubricants

Freezing Point Depression Heat transfer fluids

2.2.2 Health and environmental consideration

2.2.2.1 Health concern

According to product safety assessment documents from Dow Company, TEG itself is safe

for human health. It does not create any concern about carcinogenicity or mutagenic with

laboratory animals. However, in dermal pathway, it might be irritated when skin or eye

exposure directly with TEG. When there is an injury on skin, this exposure could be more

serious. Once massive contact with skin is sufficient, it is hot enough to burn skin. This

problem relates to absorption characteristic of TEG.

Due to the low volatility characteristic at room temperature, TEG is very safe in inhalation

exposure. In case of repeating aerosol exposure at the excessive level (high doses), it might

cause of death. In food chain, TEG’s bio-magnification is very low due to its properties.

2.2.2.2 Environmental concern

TEG is basically a non-toxic chemical. It does not harmful for aquatic microorganisms.

TEG survives in atmosphere with a very low concentration because it is very easy to be

photodegraded under sunlight (short haft-life). Due to the two important properties which

11

are soil mobility and biodegrades readily, TEG concentration in natural environment is

very low (Dow, 2014). Within 20 days, TEG can be degraded 90 % (measuring by BOD20).

2.2.3 TEG Recovery methods

In literature, amount of research on treatment of TEG wastewater is very less. However,

some authors had conducted their study on treatment and/or concentrate ethylene glycol

from wastewater. Evaporation process had proven as an effective method when it can

concentrate ethylene glycol up to 70 % (Jehle et al., 1995). Unfortunately, this is a very

slow treatment process and consume high energy. Thus, evaporation is not a really

promising process. Using nanofiltration (NF) process on concentrating ethylene glycol was

studied (Orecki et al., 2006). However, the rejection of NF membrane was failed in all tests.

2.2.3.1 Conventional membrane process

Larpkiattaworn (2013) studied about TEG removal by using polyethersulfone (PES-NTR

7450) membrane. From this study, 99 % TEG was rejected at the condition of feed

temperature 28oC and applied pressure of 1 kg/cm2.

By using two nanofiltration (NF) and two reverse osmosis (RO) membrane, Jacob (2014)

found out that the membrane’s selectivity were lost when the TEG concentration in feed

solution was higher than 10%. The author conducted the rejection test for both membranes.

The result of rejection examination of NF and RO were 80 and 83-95 % respectively. At

the initial concentration of TEG of 5%, the highest TEG concentration achieved was 89.12

and 95.74 % for RO-ACM5 and RO-NTR759 membranes respectively. The author had

concluded that membrane based treatment is effective only for wastewater that has low

initial TEG concentration (0.1-5%).

2.2.3.2 Distillation process

In the normal condition, the boiling point of water (100oC) is much lower than TEG

(288oC). Thus, by applying heat to TEG wastewater, water is first vaporized and be

collected. After dewatering step, heat is continuously applied until the temperature of waste

solution reaches the boiling point of TEG. Similar with dewatering process, TEG is distilled

and be collected. However, this process requires high energy.

2.2.3.3 Membrane distillation

Glycol separation had been studied in three configurations. The first study was conducted

with direct contact membrane distillation (DCMD) configuration by Rincón (1999).

Ethylene glycol could achieve 70% of concentration by using DCMD operating at

moderate temperature and atmospheric pressure. However, this mark was also the

limitation of DCMD. It could not achieve higher glycol concentration than that value. The

adverse effects of temperature and concentration polarization were a problem which a

careful attention must be paid to this issue. Using vacuum membrane distillation (VMD)

to concentrate ethylene glycol was studied by Mohammadi and Akbarabadi in 2005. These

authors concluded that ethylene glycol had ability to be recovered by VMD process. The

effectiveness of TEG rejection was achieved 100% in VMD configuration. However, VMD

consumes high energy than other membrane distillation configurations.

12

In membrane distillation process, it is certainly a need of an external energy source to heat

up the feed solution. Thus, optimizing energy consumption becomes attractive field for

further study.

Comparing between Distillation, NF/RO and MD

Since MD is the new technology which is not widely applied at industrial level, the

economic feasibility of MD has not been evaluated completely. To operate a MD system,

the basic standard of energy required is 628 kWh/m3 (Camacho et al., 2013) while the

energy consumption of water production required for a RO system is only 2.49 kWh/m3

(Liu et al., 2011)

Kesieme et al. (2013) conducted the study on a desalination plant that has a capacity of

30,000 m3/day with different technologies. The authors indicated that it is not economical

when comparing between MD and RO/MED if the plant is operated by supplied steam (in

this case, the production cost of MD, MED and RO are 1.72, 1.48 and 0.69 $/m3

respectively). However, the production cost of MD can reduce to 0.57 $/m3 by using waste

heat. Thus, MD would be the promising technology which against other conventional

processes.

2.3 Membrane Science and Technology

In the trend of development of water and wastewater treatment technology, membrane

technology had been developed and become important rapidly. This technology has many

potentialities to rationalize of operation process. The appearance of membrane technology

was to adapt the three important aspects: water scarcity (reclamation for water reuse),

regulatory pressure (stricter standards), and treatment cost improvement (economic

efficiency). In the field of environmental treatment, membrane technologies has been

increasingly applied exponentially within recently years (Metcaf, Eddy, 2003) and will

continue dramatically in the future.

In many places on over the world where water supplies are restricted of quantity and

quality, the concepts of reclamation, reuse and protection of water have been played a very

important role (Daigger et al., 2006) to reduce water footprint – an indicator for water reuse.

By using ultrafiltration (UF), Giardia and Cryptosporidium protozoa would be removed

(which conventional treatment process could not eliminate) completely. The strictness of

water standard (for both reuse and discharge) is rising along with timeline. For example,

nitrogen and phosphorus are required to reach the stringent standard before discharge to

reduce eutrophication phenomenon, by using membrane bioreactor (MBR), both biological

and chemical nitrogen can be removed successfully (Daigger, Crawford, 2005). From the

aspects of cost improvement, there are low cost in both membrane material (i.e. woven

fiber microfiltration (WFMF) (Thanh, Dan, 2013)) and treatment process by consuming

less energy, chemical, land use and labor while producing more water and remove more

impurities (zero discharge concept).

The main mechanism of membrane treatment is pore-filtration process, like a factitious

kidney, by providing physical barriers. From the difference in size of pores, there are

different applications and so that various materials are rejected by the pore on the surface

13

of membrane. The smaller of pore size, higher pressure of feed water needed to operate the

filtration process. Also, the properties of surface are very important. It is needed to provide

higher pressure for hydrophobic surface than hydrophilic. The typical membrane material

used for wastewater treatment is organic compound, includes: polypropylene, cellulose

acetate, aromatic, polyamides, and thin-film composites (TFC). By tailoring and adjusting,

membrane properties can be matched with any specific design of separation tank. It has

ability to upscale or connect with other treatment process to achieve higher efficiency.

However, there are also some disadvantages of membrane process, such as requirement of

chemical pretreatment or membrane fouling, and in operating process, it also can be

destroyed by incident or shock-loading. Overall characteristics of six membrane process

are shown in Table 2.5.

2.4 Membrane Distillation

In the early years of the second half of the 20th century, Membrane distillation (MD) was

first introduced as a non-equilibrium thermodynamics (thermally driven) membrane

operational process (Khayet and Matsuura, 2011; Lawson and Lloyd, 1997) The concept

of “Membrane distillation” is used to define the membrane process that has six

characteristics as following: Must be a porous structure, the liquid process does not lead

to wet the membrane, condensation must not be happened within the pores, the equilibrium

of vapor-liquid of all components must not be changed in MD process, liquid phase must

be directly exposed with one side of membrane, each partial pressure of different

components in the vapor all is the driving force of the process (Khayet and Matsuura, 2011;

Smolders and Franken, 1989).

2.4.1 Membrane distillation process principle

The MD process can be described as: feed liquid solution is firstly heated up until desired

temperature, after that it will be transferred to expose with one side of membrane

(hydrophobic surface). The volatile compounds will cross the membrane pores and reach

the inside space of fiber. The non-volatile compounds will be re-circulated to feed tank.

The hydrophobic nature of porous membrane distillation will avoid liquid from going into

the pores. Only water vapor survives inside this space. The liquid-vapor separation takes

place at the entrance of each pore.

By applying various methods at the permeate side to collect the water vapor penetrating

the porous membrane, this has resulted in different MD configurations

There are four common designs which are described clearly in Figure 2.7, includes: (1)

Direct contact membrane distillation (DCMD), (2) Vacuum membrane distillation (VMD),

(3) Air gap membrane distillation (AGMD), (4) Sweeping gas membrane distillation

(SGMD).

http://en.wikipedia.org/wiki/Non-equilibrium_thermodynamics

14

Figure 2.6 Vapor-liquid interface in membrane distillation (Lawson and Lloyd, 1997)

Figure 2.7 Four configurations of membrane distillation process

2.4.2 Direct contact membrane distillation (DCMD)

Warm Feed

Cold

liquid

Porous Membrane

a. Direct contact membrane distillation

Warm Feed

Sweeping

Gas

Porous Membrane

b. Sweep gas membrane distillation

Warm Feed

Applied

Vacuum

Porous Membrane

Permeate out

Warm Feed

Porous Membrane

Air

Gap

Coolan

t

d. Air gap membrane distillation c. Vacuum membrane distillation

Aqueous

Solution

θ

Aqueous

Solution,

Sweep

Gas,

Vacuum

Or

Air Gap

15

In DCMD, condensing liquid solution with low temperature is used at the permeate side.

Both feed aqueous solution and permeate liquid solution are directly expose to the

membrane surface. The difference in temperature of the two liquid flows is the driving

force of this process. As a result, there are two interface at both sides of each pore, the

interface of vapor/hot aqueous at the entrance and vapor/cold at the permeate side. The

partial pressure of cold liquid at permeate side can be reduced to enhance the driving force

of DCMD by using osmosis distillation (OD) water (Laganà et al., 2000)

This configuration of MD is the most popular application to carry out the experiment

(Andrjesdóttir et al., 2013; Khayet and Matsuura, 2011; Phattaranawik and Jiraratananon,

2001; Qtaishat et al., 2008). For the separation of heated influent flow is water (desalination

or concentration of liquid solution), DCMD is the best solution (Laganà et al., 2000;

Lawson and Lloyd, 1997).

2.4.3 Vacuum membrane distillation (VMD)

In VMD configuration, a vacuum pump is used to maintain vacuum condition at the

downstream side of the membrane. The saturation pressure of volatile material that needed

to be separated must be higher than vacuum pressure (Khayet and Matsuura, 2011; Lawson

and Lloyd, 1997). The driving force of VMD processing is the difference in the pressure

between two sides or each pore. Comparing the permeate flux of RO and four

configurations of MD, VMD provided highest value. This result is achieved due to reducing

downstream pressure (Bandini et al., 1992; Izquierdo-Gil, Jonsson, 2003). However, the

main adverse point of VMD is the great opportunity of pore wetting due to the negative

pressure at the permeate side of membrane. Hence, VMD will operate better with the

smaller pore size (Khayet and Matsuura, 2011; Lawson and Lloyd, 1997).

The applications of VMD is described clearly by (Sarbatly and Chiam, 2013), includes:

Desalination (NaCl/water), concentration (LiBr/water), extraction of trace volatile organic

compounds (ethanol/water), removal of dissolved gases (Ammonia/water), preservation of

aroma compounds (Must/aromas/water), recovery of aroma compounds (blackcurrant

aromas).

2.4.4 Air gap membrane distillation (AGMD)

In AGMD, only one side of MD is exposed with feed heated aqueous solution, the

remaining side is unattached (Meindersma et al., 2006). Between the permeate side of

membrane surface and the surface of condensation, a stagnant air gap is located. Before

condensing at the cold condenser, the vapor of volatile material has already been passed

both the porous structure membrane and air gap (El-Bourawi et al., 2006; Khayet and

Matsuura, 2011). So, AGMD produced a lowest permeate flux in comparison with other

configurations of MD.

16

Table 2.5 General Characteristics of Membrane Process (Metcalf and Eddy, 2004)

Membrane

process

Membrane

driving force

Typical

separation

mechanism

Operating

structure

(pore size)

Typical

operating

range (µm)

Permeate

description

Typical constituents

removed

Microfiltration

(MF)

Hydrostatic

pressure difference

of vacuum in open

vessels

Sieve Macropores

(>50 nm) 0.08-2.0

Water,

dissolved solutes

TSS, turbidity,

protozoan oocysts and

cysts, some bacteria and

viruses

Ultrafiltration

(UF)

Hydrostatic

pressure difference Sieve

Mesopores

(2-50 nm) 0.005-0.2

Water; small

molecules

Macromolecules,

colloids, most bacteria,

some viruses, proteins

Nano filtration

(NF)

Hydrostatic

pressure difference

Sieve + solution/

diffusion +

exclusion

Micropores

(<2nm) 0.001-0.01

Water, very

small molecules,

ionic solutes

Small molecules, some

hardness, viruses

Reverse osmosis

(RO)

Hydrostatic

pressure difference

Solution /

diffusion +

exclusion

Dense

(<2nm) 0.0001-0.001

Water, very

small molecules,

ionic solutes

Very small molecules,

color, hardness, sulfate,

nitrate, sodium, other

ions

Dialysis Concentration

difference Diffusion

Mesopores

(2-50 nm) -

Water; small

molecules

Macromolecules,

colloids, most bacteria,

some viruses, proteins

Electro- dialysis

Electromotive force

Ion exchange

with selective

membranes

Micropores

(<2nm) -

Water, ionic

solutes Ionized salt ions

17

The gas between membrane surface and cold surface is a barrier that resulted in reduction

of head loss (Meindersma et al., 2006). The vapor flux has to be maintained to overcome

the air gap barrier. This flux is affected by the width of air gap.

Air gap membrane distillation is suitable for seperation of alcohols/liquid solution (Garcı́a-

Payo et al., 2000). This separation could not be implemented by DCMD because volatiled

alcohol is probable to wet the pore at permeate side because of lower surface tension

(Meindersma et al., 2006).

2.4.5 Sweeping gas membrane distillation (SGMD)

SGMD has another name as membrane air stripping (Meindersma et al., 2006). Similar

with VMD, vapor is condensed at the place outside membrane module. The mechanism of

this process operation is the removal of vapor by a sweep gas at the permeate side of

membrane. In SGMD, the advantage is the small resistance of air barrier that affected to

mass transfer. However, the vapor will be diluted in sweep gas that resulted in requirement

of higher condenser capacity. Furthermore, sweeping gas is easily and fast heated up by

the temperature from the vapor. Consequently, the vapor pressure would be increased to

higher level which has resulted in reducing the driving force of this operating process. In

this MD configuration, the flux is not depend of the temperature of the gas (Lawson, Lloyd,

1997). Similar with AGMD, SGMD is mostly used to remove volatile compound more than

water (Khayet and Matsuura, 2011; Zhang et al., 2000).

2.5 Material of Membrane and Module Fabrications Using for Distillation Process

The selection of MD is depended on each typical required application. It is a combination

of permeate flux, thermal conductivity, pore size, porosity, separation factor (Khayet,

2011).

2.5.1 Commercial membranes used in distillation process

To prevent the membrane wetting phenomena, hydrophobic polymer is an appropriate

material for this micro-porous membrane. Many different polymers could be used, such as:

Polypropylen, Polyvinylidene fluoride, polythylene, polytetrafluoroethylene with the

abbreviated forms as PP, PVDF, PE, PTEE respectively. These material are survived in

various shapes, such as: tubular, capillary, flat sheet. The morphological configurations of

these synthetic materials are close to meet all requirements of MD process.

2.5.2 Fabricated membranes for distillation Process

Based on different materials, various hydrophobic porous membranes are created by using

diverse techniques. The choice and production are relied on various factors, includes:

aqueous solution, range of temperature that MD can operate with, thermal conductivity,

price, easy or difficult to fabricate and assembly. Flat sheet, hollow fiber with single

hydrophobic layer membrane and composite multilayers membrane (bi-layer of

hydrophobic/hydrophilic membrane) are created and used (Khayet, 2011).

2.5.2.1 Flat sheet and frame module

18

Many types of flat sheet membranes with single hydrophobic layer have been developed

and applied for MD processing, such as: asymmetric PVDF (polymer concentration of 10-

25wt.%, solvent used of 13.2 -15 wt.% , porosity >79%, pore size of 0.0698-0.349 µm),

copolymer PVDF-TFE (pore size < 2.4 x10-2 µm, porosity < 80 %), copolymer PVDF-HFP

(19.1 wt.% PVDF-HFP, solvent polyethylene glycol (PEG) of 4.99 wt.%) (Khayet and

Matsuura, 2011).

Depends on each membrane that has different concentration of polymer and using different

amount of solvent, there will be the difference in the maximum coefficient of mass transfer

of different membranes. This module is generally used in laboratory studies because it can

be cleaned and replaced easily. Notwithstanding, it is very low in the value of the ratio

between area of membrane to the module volume.

2.5.2.2 Hollow fiber module

By using different polymers, different solvents and different spinning process (dry/wet),

the different hollow fiber membrane were made out. Some common materials for hollow

fiber are: PVDF, PVDF/Cloisite clay, PTFE, copolymers (Khayet, 2011). An typical

example, PVDF membrane material with pore size of 4.0-24.8 nm, porosity of 56-73 %,

internal diameter of 0.675-0.844 mm, external diameter of 0.982-1.071 mm (Fujii et al.,

1992).

The main composition of this module is a shell tube that includes a determined number of

hollow fibers bundled and sealed. The significant advantages of this type of module are

low energy consumption and high membrane area in a limited volume. Contrarily, it is

difficult to clean and high opportunity to get fouling.

2.5.2.3 Spiral wound module

The components of this type include a flat sheet membrane that is rolled and enveloped in

a limited space. The center of the winding is a collection pipe. The movement of feed

solution overpass the membrane surface is in an axial trend while the permeate flux goes

into the central tube. Alkhudhiri (2012) confirmed that this module has high packing

density, not easy to fouling, and the energy consumption is acceptable.

2.6 Operational Processes of Membrane Distillation

2.6.1 Development of theoretical models for membrane distillation

One side of membrane must be directly contact with the feed heated aqueous solution. The

membrane aqueous solution entry pressure (liquid entry pressure – LEP) must be higher

than the applied hydrostatic trans-membrane pressure. The surface tension force of

membrane (hydrophobic nature) prevents liquid from entering the pores. Generally, there

is a supposition of the negligible kinetic effect of liquid/vapor interface and the equilibrium

of liquid/vapor phases is directly corresponding to the temperature at membrane surfaces

when developing a MD model process (Khayet, 2011).

The formula is used to calculate separation factor for the feed solution containing non-

volatile materials is as following:

19

α = (1 −Cp

Cf) 100 (2.5)

Where α is the separator factor, 𝐶𝑝 is the solute concentration in the permeate flux and 𝐶𝑓

is the concentration in feed flow.

If the volatile compounds are contained in feed solutions, the above formula will be

changed to:

α =Xv,p/Xw,p

Xv,f/Xw,f (2.6)

With X𝑣,𝑝, X𝑤,𝑝are the mole fractions of volatile compounds (v), water (w) in the permeate

(p) flux and X𝑤,𝑝, X𝑤,𝑓 are those values in feed (f) solutions.

The vapor pressure of a given compound (i) is calculated by Antoine Equation as following:

pi(T) = exp (α −β

γ+T) (2.7)

Where pi and T are the partial vapor pressure of the pure component in the permeate flux

(Pa) and absolute temperature (K), α, β, γ are constants that depend on typical material. i

can be water or any chemical compound (for water, α = 23.1964, β =3816.44, γ = - 46.13).

By calculating the condensate collected in the permeate side of MD module for a

determined time, the permeate flux (Ji) in all MD configurations (that depends on the nature

characteristics of membrane and driving force) would be measured

Jw = Bw∆pw = Bw(p𝑚,𝑓 − p𝑚,𝑝) (2.8)

Where J is kg/m2h, , Bw is the membrane coefficient (permeability of MD), p𝑚,𝑓 is the

partial vapor pressure of water in the feed solution, p𝑚,𝑝 is its value in permeate side.

If the feed solution is diluted aqueous of non-volatile materials, the partial vapor pressure

of that solution can be calculated as:

pw,s = (1 − xs)pw (2.9)

With pw,s is the vapor pressure of a diluted aqueous of non-volatile materials, pw is the

vapor pressure of water, xs is the mole fraction of non-volatile compound.

Liquid entry pressure

The concept of liquid entry pressure (LEP) is used to prevent the wetting phenomenon of

membrane pores. LED is the lowest pressure that needs to be applied for the feed aqueous

solution before touching the entrance of the dry pores. The value of LEP can be measured

by using the Laplace Equation (Lawson, Lloyd, 1997) which is expressed as the following:

20

LEP > ∆Pinterfae = Pliquid − Pvapor =−2BγL cos θ

rmax (2.10)

Where , B, γL, θ, rmax are geometric coefficient which is measured by pore structure,

aqueous surface tension, the angle of the contact between liquid and solid (feed aqueous

solution and the surface of membrane), and the largest pore size respectively. In general,

VMD uses the membrane which has rmax < 0.45 µ𝑚 (Lawson, Lloyd, 1997). The value

of LEP reduces along with the increase of rmax of membrane and/or the reduction of θ.

2.6.2 Mass transfer in membrane distillation

There are two contingents of mass transfer through MD. One part is volatile compounds in

vapor will cross the pores of MD then it will pass the boundary layers of membrane surface

in the second part. This layer is a thin film between membrane surface and bulk aqueous

solution and will be discussed in details in the polarization section.

The permeate flux which is through the micro-porous membrane can be anticipated exactly

by using dusty gas model. This model is the combination of four componential

mechanisms: Knudsen diffusion, ordinary molecular diffusion, viscous/poiseuille of flow

(surface diffusion is neglected in dusty model) (Khayet and Matsuura, 2011). The typical

expression is shown in Figure 2.8.

Dusty gas model

In water, the solubility of air is about 10 ppm (Khayet, 2011). Generally, mass transfer in

MD is the results of convective and diffusion of volatile material which cross the pores of

membrane. Knudsen diffusion model and Viscous/Poiseuille model are used to describe

the resistance of micro-porous structure membrane in absence of air. Molecular diffusion

model is used to describe the mass flux in presence of air. In dusty gas model that using for

DCMD, surface diffusion is considered as neglect due to the very small membrane surface

in comparison with the total area of pores. Besides, the operating pressure of DCMD is

always maintained at a constant value (~105 Pa) and the flow of vapor that passing the

membrane porous is very small relative to the water flux, thus viscous flow is not

considered as significantly negligible (Lawson and Lloyd, 1997).

Knudsen number (Kn) is the quantifiable value which is used to determine the operational

mechanism of a pore of MD under a typical condition. This number is calculated as a ratio

of the mean free path (λ) of a given compound (the mean free path can be defined as the

average route of a moving molecule between each successful collision)n. This collision

must change directly energy or direction of that molecule) to the pore size MD (dp).

Kn =λ

dp (2.11)

The value of mean free path of a given molecules (λi, m) can be measured by using the

following Equation:

λi =Tkb

(√2)Pm(σi)2 (2.12)

21

Where σi is the collision diameter of given specie (for water molecules in gas phase, σw =2.641 × 10−10 m), kb is Boltzmann constant, Pm is the mean pressure within the pores,

and T is the absolute temperature.

Figure 2.8 Dusty gas model

Membrane distillation coefficient is calculated by Equation 2.13.

Bik =

2πrk3

3RTτδ (

8RT

πMi)

12⁄

(2.13)

Where 𝑟𝑘, 𝑀𝑖, 𝛿, 𝜏, 𝑅 are pore radius of membrane, molecular weight of given specie,

thickness of membrane, tortuosity of membrane and gas constant respectively.

In case Kn < 0.1 (dp > 10λi) , molecular-molecular impinge (molecular diffusion) is the

main responsibility for mass transfer the pressure of MD operating system in this situation

is nearly approximate to atmospheric pressure (Qtaishat et al., 2008). The Equation 2.14

below expresses the membrane coefficient of MD that has a pore’s area of 𝜋𝑟𝐷2 in the

region of ordinary diffusion model.

BiD =

πPDirD2

RTPaτδ (2.14)

Where Pa is the partial pressure of vapor inside the pore, P is total pressure within the

membrane pore, Di is the diffusion coefficient of a given material.

For air/water solutions, PDi can be measured by the Equation below where PDw value is

expressed in the unit of Pa m2 s⁄

PDw = 1.895 × 10−5 × T2.072 (2.15)

If the value of Kn is in the range from 0.01 to 1 (100λi > dp > λi), the transportation

mechanisms of given volatile compounds are molecule-wall and molecule-molecule

diffusion (Knudsen and ordinary diffusion) which are taking place in combination. Micro-

porous membrane’s permeability of the pores that has an area of 𝜋𝑟𝑡2 can be calculated by

using Equation 2.16.

Knudsen Ordinary diffusion

Rv = 0

Viscous

Rs = 0

Surface

22

BiC =

π

RTτδ[(

2

3rt

3 (8RT

πMi)

12⁄

)

−1

+ (PDirt

2

Pa)

−1

]

−1

(2.16)

Khayet (2011) indicated that the combined Knudsen/molecular diffusion mechanism is

dominant for the membranes which has pore sizes in the range from 0.2 to 1 µm. When the

mean free path is close to to the mean pore size, permeate flux will not increase along with

the opening of pore size. Consequently, MD will work better with the membrane that has

smaller pore size than the mean free path (achieving higher flux under Knudsen diffusion

mechanism)

Khayet and Matsuura (2011) found that the mechanism responds for water vapor across

the pores of membrane is the combination of Knudsen diffusion and molecular diffusive

flux. The total permeate flux can be presaged exactly by theoretical model, but it is over

ability to estimate the partial organic permeate flux.

The partial pressure of water vapor (pw,p) in the permeate flux at the permeate side can be

calculated by using the formula as below:

pw,p =P.w

w+0.622 (2.17)

Where P is the total pressure of permeate flux at the permeate side, w is the humidity ratio.

The value of pw,p is depended on the gas temperature at the surface of micro-porous

membrane.

The humidity ratio is determined for a typical given air sample, can be understood as a

portion of mass between water vapor and dry air. The value of w can be achieved by the

relation as following:

w = win +AJw

ma (2.18)

In the Equation above, win is the ratio of humidity at the inlet point of module, ma is the

air flow rate, A is membrane effective area, Jw is the total permeate flux achieved in SGMD

configuration.

The second-level formula for the permeate flux of water vapor is the combination of some

equations that presented as above.

Jw2 + Jwb + c = 0 (2.19)

The value of b and c (coefficients) are estimated from the following respectively:

b = Bw(P − awpw,f0 ) +

ma

A(win + 0.622) (2.20)

c = Bw

Ama (Pwin − awpw,f

0 (win + 0.622)) (2.21)

23

Where 𝑝𝑤,𝑓0 is partial pressure of pure water in permeate flux, aw is the activity of water,

𝐵𝑤 is the coefficient of SGMD (permeability, productivity).

If there is a mixture of specie i and j in the permeate vapor through the membrane pores,

the mean free path can be measured by the Equation as following:

λi/j =TKB

πPm(√1+Mi Mj⁄ )((σi+σj) 2⁄ )2 (2.22)

Where σi, σj and Mi, Mj are collision diameters and molecular weight of species i and j

(volatile compounds) respectively.

Once the mean free path of given molecules is shorter than the pore size of MD, molecule-

molecule collisions will become the main phenomenon for mass transfer, over the

molecule-wall collisions. Poiseuille (viscous) type of flow is the main mechanism in this

case. Consequently, Bi is evaluated by Equation 2.23:

Bi =εr2Pm

8τδniRT (2.23)

Where 𝑛𝑖, 휀 and P are viscosity of typical materials, porosity of membrane, and average

hydrostatic pressure within the pores respectively.

Figure 2.9 Poiseuille type of flow inside a pore of SGMD process

2.6.3 Heat transfer in membrane distillation

Membrane distillation process is operated by the combination of two processes: mass and

heat transfer process, which are happened simultaneously. The heat transfer in MD process

can be separated in three steps: (i) heat transfer cross the boundary layer at the feed side of

membrane surface, (ii) heat transfer throughout the pores of micro-porous membrane, (iii)

heat transfer cross the boundary layer at the permeate side of membrane surface. Figure

below divided total heat flux (𝑄𝑚) into two heat transfer mechanisms: (i) heat transfer

24

through membrane material (membrane wall) and heat of pores that filled up by gas (Qc)

and (ii) heat of volatile molecules in vapor flux (Qv).

Heat transfer at the two boundary layers which are feed side (Qf) and permeate side (Qp)

of membrane surfaces respectively as the following:

Qf = hf(Tb,f − Tm,f) (2.24)

Qp = hp(Tm,p − Tb,p) (2.25)

Where hf and hp are coefficients of heat transport though two boundary layers which are

mentioned above. The acronyms Tb,f, Tm,f, Tm,p, Tb,p in two equations above are

temperatures at bulk feed, membrane-feed (membrane-solution) interface, membrane-

permeate interface, bulk-permeate (vapor-liquid) interface respectively.

For the heat flux that transferring through membrane material (Qc) and heat trapped in the

vaporized molecules (Qv) (Qv can be understood by heat transfer accompanies with mass

transfer), its balance is performed as following:

Qm = Qc + Qv (2.26)

Two partial heat transfers above are expressed as the two following Equations:

Qv = Ji∆Hv,i (2.27)

Qc =km

δ(Tm,f − Tm,p) = hm(Tm,f − Tm,p) (2.28)

Where Ji, ∆Hv,i, km, hm, δ are permeate flux, latent heat of vapor molecules of specie i,

thermal conductivity of micro-porous MD, heat transfer coefficient of the whole

membrane, membrane thickness respectively.

The heat transfer coefficient of only vapor flux (hv) is measured as:

hv =Ji∆Hv,i

(Tm,f−Tm,p) (2.29)

Khayet (2011) indicated that 50-80 % of energy consumption is accounted for Qv, the

residual part is for Qc and Qc is considered as head loss. The efficiency of heat transfer (η)

is calculated by the Equation below:

η =Qv

Qc+Qv=

Ji∆Hv,i

Ji∆Hv,i+hm(Tm,f−Tm,p) (2.30)

In the stable condition:

Qf = Qm = Qp = Q (2.31)

25

So, the heat transfer process in MD (boundary layer at feet side – membrane material –

boundary layer at permeate side) can be expressed in the summary in the following

Equation:

hf(Tb,f − Tm,f) =km

δ(Tm,f − Tm,p) + Ji∆Hv,i = hp(Tm,p − Tb,p) = hc(Tb,f − Tb,p)

(2.31)

And

Q = ∆T [1

hf+

1

hp+

1

km δ⁄ +Ji∆Hv,i ∆Tm⁄]

−1

= ∆Thc (2.32)

With

hc = [1

hf+

1

hp+

1

km δ⁄ +Ji∆Hv,i ∆Tm⁄]

−1

(2.33)

Where ℎ𝑐 is the heat transfer coefficient of the whole MD process, ∆T (∆T = Tb,f − Tb,p)

is the bulk temperature disparity between feed aqueous solution and permeate flux, and

∆Tm (∆Tm = Tm,f − Tm,p) is the disparity of transmembrane temperature.

2.7 Temperature Polarization

Boundary layers are the limiting barriers of MD efficiency in the heat transfer process. To

quantify the size of partial resistance of the boundary layers over the total resistance of

whole heat transfer process, temperature polarization coefficient (TPC) is usually used.

TPC performs the driving force reduction (∆pi) and is calculated by the following

Equation:

TPC =Tm,f−Tm,p

Tb,f−Tb,p=

1

1+hc h⁄= 1 − hc h⁄ (2.34)

Where

h = (1

hf+

1

hp)

−1

(2.35)

According to Khayet (2011), TPC indicates whether MD design is good or not good. In

case of TPC smaller than 0.2, MD process is limited in heat transfer and very poor in

module design. If TPC is higher than 0.6, MD process is limited in mass transfer and poor

in productivity (low permeability).

If TPC is approximately close to 1, the heat transfer across both layers achieves a high

efficiency and the effect of thermal polarization is neglected and MD process is controlled

by resistance of mass transfer. If TPC is close to 0, the efficiency of heat transfer through

both layers is very low, the effect of thermal polarization is very high and MD process is

controlled by resistance of heat transfer. The desired MD system has the TPC value in the

range from 0.4 to 0.7. In other words, the difference of temperature at permeate boundary

layer from feed aqueous boundary layer is between 30% and 60%. TPC value can be

increased by increasing feed aqueous solution and permeate flow rate and decreases

temperature of feed solution.

26

TPC can be understood by the relation in the following Equations:

TPC = TPCf + TPCp − 1 (2.36)

TPCf = 1 −hc

hf=

Tm,f−Tb,p

Tb,f−Tb,p (2.37)

TPCp = 1 −hc

hp=

Tb,f−Tm,p

Tb,f−Tb,p (2.38)

Where TPCf and TPCp are heat polarization coefficients of feed and permeate flux. The

value of hf is always higher than hp (Khayet et al., 2002).

Figure 2.10 Temperature and concentration profile in MD process

2.8 Concentration polarization

In MD operational process, volatile solutes pass the membrane pores and micro-porous

membrane rejects non-volatile material, concentrating it on the membrane surface at the

feed side. The concentration of non-volatile compounds at the interface of membrane-feed

aqueous is higher than that in the bulk solution. Consequently, driving force of MD process

and permeate flux are reduced.

The value of concentration polarization coefficient (CPCi) of a given non-volatile

compound appear in feed solution is calculated as:

CPCi =Cm,f

Cb,f= exp (Jw ks⁄ ) (2.39)

Temperature

Concentration

Membrane

Boundary layer Boundary layer

Permeate Feed

27

Where Cm,f, Cb,f are concentrate of solute at the feed-membrane film and in the bulk feed

aqueous solution respectively, ks is coefficient of mass transfer. The value of ks across

the boundary layer can be calculated by using Equation below, with D, δ are molecular

diffusivity and thickness of the boundary layer respectively.

ks =D

δ (2.40)

In all MD configurations, the contribution of concentration polarization of boundary layers

in total resistance of mass transfer is less than temperature polarization (Khayet et al.,

2004).

As mentioned previously, the operational driving force of MD process is reduce because

of both temperature and concentration polarization. Hence, it can be become simple when

both TPC and CPC are expressed in one specific coefficient, vapor pressure polarization

coefficient (Ψ). Ψ is typically understood as the portion of externally driving force that

motivates the mass transfer.

Ψ =pm,f−pm,p

pb,f−pb,p (2.41)

Where pb,f, pm,f, pm,p, pb,p are partial pressure of a given specie at bulk feed, membrane-

feed (membrane-solution) interface, membrane-permeate interface, bulk-permeate (vapor-

liquid) interface respectively.

Figure 2.11 Concentration and temperature polarization in MD process

2.9 Membrane Fouling

In MD process, the fouling phenomenon is determined as the process of building up or

attaching of undesirable components on the surface and pores of membrane. This

phenomenon leads to decrease the efficiency of MD process, typically reducing in

permeability and approaches to pore wetting by increasing membrane resistance. The

undesirable components can be suspended solids, corrosion particles, biological growth, or

28

crystalline deposits. Once the fouling layer appears on the membrane surface, the thermal

resistance is also created. It leads to change the overall coefficient of heat transfer.

Fouling in MD process is not serious as others membrane separation process and it is easy

to overcome by some simple pre-treatment steps (Srisurichan et al., 2005). Alklaibi and

Lior (2005) concluded that the permeate flux can be enhanced up to 25% if the pre-

treatment process is implemented well. In another hand, this operational problem also can

be overcome by increasing the flow rate of feed solution and decrease the feed temperature

(Alkhudhiri et al., 2012).

2.9.1 Biological growth fouling

Some microorganisms grow and attach to the surface of membrane. The microorganisms

can be algae, bacteria, seaweed and fungi. Once this phenomenon takes place, a

significantly decline in permeate flux is happened together because of scale formation.

Biological growth can be reduced or prevented by boiling feed solution about half an hour

then filtrate this solution by appropriate filtration technology (El-Bourawi et al., 2006)

Figure 2.12 Profile of temperature in fouling case

2.9.2 Suspended particles and corrosion products fouling

The accumulation of suspended particles on the membrane surface is determined as

particulate fouling. There is always an approach of particles accumulation on the membrane

surface once the feed solution travels in MD system as streamline formation and the

random motion of particles leads to reach the gradient of undesirable component

concentration. Brownian diffusion affects directly on the travelling process of a very small

particles while the momentum forces control the large particles. Moreover, membrane

pores can be plugged by the particles that have sources from the corrosion of the equipment

Temperature

Thermal

Boundary layer

Fouling Layer

Membrane

Thermal

Boundary layer

29

such as pipe line, pump or heat exchangers. It is not able to predict when the fouling

phenomenon takes place because it is a time dependent process

2.9.3 Crystallization fouling

During the desalination process, the major operational problem of this MD application is

membrane surface is scaled by crystalline deposits formation, typically salt molecules. This

phenomenon leads to reduce the hydrophobicity of membrane and the pores can be clogged

by water. Karakulski (2002) found that the permeate flux is decreased dramatically when

CaCO3 accumulates on the membrane surface during the study about demineralization in

tap water. It is worth quoting here that crystal fouling formation is undetected in AGMD

module even this system was operated in a long period ( almost 2 month) (Banat, Simandl,

1994). This fouling is generally observed in DCMD process (Tun et al., 2005). It is mainly

related to salt solubility characteristics (which is usually decreased along with the reduction

of temperature).

The hydrodynamic condition and the interaction between membrane surface and foulants

were the most important factors affecting on membrane fouling. There were three types of

fouling: organic, inorganic, and biological fouling. In organic fouling, proteins tend to

attach on membrane hydrophobic surface. This has resulted in decreasing water contact

angle on membrane surface (membrane hydrophilization). In inorganic fouling

(crystallization fouling), it is not a significant problem due to the turbulent condition of

feed flow. About the biological fouling, because of the severe conditions of feed solution

in MD process (temperature and flow rate), biofilm could not be formed on the membrane

surface. Thus, the contribution of biofouling into the total membrane fouling is same with

inorganic fouling.

2.10 Membrane Cleaning

There are two adverse effects of membrane fouling: decreasing in permeate flux and

reducing membrane selectivity (due to reducing in hydrophobic nature of membrane). The

purposes of membrane cleaning process are to remove membrane fouling layers and

recover membrane properties. By applying an appropriate pre-treatment process,

membrane fouling can be reduced. Furthermore, cleaning of MD module can help to

restore membrane properties till its initial level. Reversible fouling can be removed by

clean membrane with pure water. Chemical cleaning can help to eliminate irreversible

fouling.

Guillen-Burrieza (2014) had conducted a study about various MD cleaning strategies in

term of permeate flux recovery, quality of distillation and membrane damage. There was a

difference in chemical solution used in this research. The author found out that the most

effective chemical using to clean MD membrane were the mixture of 0.1 wt.% Oxalic acid

and 0.8 wt.% Citric acid. The cleaning procedure using in this study were: 6 h with

chemical, 30 min with pure water and 24 h of dry-out.

2.11 Operating Variables Affecting MD Process

Variable factors affecting on the MD processing in each MD configuration are shown in

Table 2.6.

30

2.11.1 Effects of the temperature of feed solution

In general, the feed liquid solution is heated up till the range from 20 to 80oC (under the

level of boiling point) while all other parameters of MD are kept constant. From Table 2.6,

it is observed that accompanying with the increasing temperature of feed solution, MD flux

also increase exponentially.

This phenomenon can be explained by the increase of vapor pressure in aqueous solution

exponentially (which leads to increase driving force of MD process) with increasing

temperature. It is noticed that MD system is operating better when working under high

temperature of feed solution. Although the thermal polarization effect becomes higher

when increasing feed temperature but the contribution of feed temperature to evaporation

and total heat transfer is very high (El-Bourawi et al., 2006; J. Phattaranawik et al., 2003).

2.11.2 Effects of feed concentration

The effects of concentration of a typical solute in feed liquid on vapor flux strongly depend

on the separation process of that self-compound. MD process can be applied to treat feed

solution which is very high concentration in comparison with other conventional membrane

process (which is operated by pressure-driven) (Banat, Simandl, 1994).

While considering non-volatile compounds in feed solution, it was found that the main

effect of itself is the reducing in vapor flux of MD processing. Consequently, it was

concluded that the driving force of MD separation process can be reduced by adding non-

volatile compound to feed solution (by decreasing partial vapor pressure). The effects of

concentration polarization also get the contribution from that (increases the thickness of

boundary layer on feed side of MD) (Tomaszewska et al., 1998)

When volatile solutes in feed solution (i.e. alcohols, ammonia) are considered, it was found

that the permeate flux would be increased when rising the volatile components

concentration in feed solution (El-Bourawi et al., 2006). Because while increasing

concentration of volatile compounds in feed solution, the partial vapor pressure of that

component in feed liquid (which plays as driving force of MD process) is also increasing.

In this situation, some cares need to be put to prevent the pore wetting phenomenon.

2.11.3 Effects of the circulation velocity and stirring rate of feed solution

The circulation and stirring are applied to achieve a higher permeate flux throughout

increasing the coefficient of heat transfer and reduces the polarization of thermal and

concentration. Operating with turbulent flow (which is achieved by rising circulation

velocity or mixing power) leads to get higher productivity of MD system (Li et al., 2003;

Yang et al., 2012). It means that the temperature disparity at two sides of the pores is higher

because the temperature of boundary layer at the feed membrane surface approach closer

to the temperature of bulk feed solution. The general effects of circulation velocity are

shown in Table 2.6.

However, the flow rate of feed solution which contains non-volatile compounds in SGMD

processing is practically considered as negligible (Khayet et al., 2000). It means that the

31

feed flow velocity affects directly on MD separation process when the feed aqueous

solution contains volatile compounds. On other words, flow rate of feed liquid solution has

a strong relationship with LEP in preventing the pore wetting phenomenon.

The turbulent condition can be expressed by the linear relationship among Reynolds

number and feed flow velocity, as the Equation below

ReN =vρdh

μ (2.42)

Where ReN is Reynolds number, v is feed flow velocity, ρ is feed liquid density, µ is

aqueous viscosity, dh is diameter of feed liquid channel.

2.11.4 Effects of the temperature of permeate inlet flux

The permeate flux is reducing while increasing the temperature of permeate inlet flux due

to the decrease of vapor pressure of the membrane process. Consequently, to achieve a

higher permeate flux, it’s more suitable to increase the temperature of feed solution than

reduce the temperature at the permeate side (El-Bourawi et al., 2006). In DCMD, the

permeate flux increase while decreasing the inlet permeate temperature (Martı́nez et al.,

2003). However, the increase of permeate temperature in AGMD configuration is

considered as zero because the heat transfer coefficient of the air gap plays as the main role

for total heat transfer coefficient, hence the is a very little effect from the changes in the

temperature of permeate flux (Banat, Simandl, 1994). It is similar to SGMD processing

due to the very fast increasing of the temperature of sweeping gas, so that changing

temperature effect from permeate flux is negligible (Lawson, Lloyd, 1997).

2.11.5 Effects of temperature difference

Temperature difference between both sides of transmembrane creates the vapor pressure

which is considered as the driving force of MD process. The vapor flux accretes linearly

accompanying with temperature difference when the mean temperature are kept as constant

(Lawson, Lloyd, 1997). It means that once the temperature difference is constant, the vapor

flux will expand exponentially accompanying with the mean temperature (Khayet et al.,

2002). In contrast, on SGMD system, it is monotonically increase instead of linearly

increase (Khayet et al., 2000).

32

Table 2.6 Effect of Variables on Permeate Flux in MD Process (El-Bourawi et al., 2006)

Configuration

Variable

Feed side operating variables, effect of increasing of Permeate side operating variables, effect of

increasing of Vapor

pressure

differnece Temperature Concentration Velocity Stirring rate Temperature Velocity Stirring rate

DCMD

AGMD

SGMD

VMD

Remark: : increasing with; : decreasing with; : not clear;

Table 2.7 Effect of Membrane Parameter on Permeate Flux in MD Process (Khayet and Matsuura, 2011)

Remark: : increasing with; : decreasing with; : not clear; : slightly effect.

Configuration Parameter

Thickness Porosity Pore size Pore size distribution Tortuosity

DCMD

AGMD

SGMD

VMD

33

This phenomenon can be explained as: if the temperature of vapor flux is maintained

constant while rising the temperature of feed solution, the permeate flux will increase

monotonically with the temperature difference.

2.11.6 Effects of the velocity of permeate flow

Only for DCMD and SGMD, the permeate flow effects can be observed. If the permeate

flux contains volatile compounds, increasing velocity of permeate flow leads to decreasing

the effect of thermal polarization and concentration polarization. Consequently, the heat

transfer at permeate side would be increased.

Once the value of coefficient of heat transfer at the permeate side gains, the temperature of

the bulk permeate flow is closer to the temperature of the membrane permeate surface, as

a result, the driving force of MD process would be increased (El-Bourawi et al., 2006). In

SGMD configuration, the thermal polarization is appeared at the membrane permeate side

when the feed solution contains non-volatile compounds (Khayet et al., 2000). By

decreasing permeate flux or increases permeate velocity, we can achieve an optimum of

permeate flow velocity. Hence, the highest transmembrane flux can be achieved by

optimizing the air pressure flow in the membrane permeate side. In general, the feed side

pressure must be higher than the pressure at permeate side to operate MD process and LEP

of feed liquid solution must be higher than transmembrane hydrostatic pressure to prevent

the pore wetting phenomenon.

2.11.7 Effects of vapor pressure difference

The driving force of MD process which is understood by vapor pressure difference between

two sides of membrane can be implemented by temperature difference or applies a vacuum

pressure at the membrane permeate side. It is proved that permeate flux linearly increases

with driving force in all MD process generally. In fact, it is not clear in the effect of

transmembrane hydrostatic pressure, so it is necessary to study more about this in all MD

processes (El-Bourawi et al., 2006).

2.12 Effects of Membrane Parameters on MD Process.

In fact, most of materials using for MD process are produced from polyvinylidene fluoride

(PVDF), polytetrafluoroethylene (PTFE) and polypropylene (PP). A satisfy micro-porous

membrane must has low mass transfer resistance, high LEP to keep pore dry, high stability

of thermal, high resistance of chemical in liquid solution. The interface between variable

membrane characteristics and permeate flux is expressed by an Equation below (Lawson,

Lloyd, 1997):

𝑀 ∝𝑟𝑛

𝛿𝜏 (2.42)

Where M is molar flux through a typical pore of membrane which is proportional to the

ratio between the multiplication of membrane porosity (휀) with mean pore size of

membrane (n = 1 for Knudsen diffusion, n = 2 for viscous diffusion) and the multiplication

of membrane thickness (𝛿) with membrane tortuosity (𝜏). Table 2.7 shows the effect of

each typical membrane characteristic for all MD processes.

2.12.1 Effects of membrane thickness

34

The thickness of membrane affects directly on mass transfer process via the value of

resistance. The less thin as possible of membrane, the higher membrane permeability is

achieved (El-Bourawi et al., 2006) . Contrarily, because heat loss in membrane module is

implemented at the membrane material, so the better heat efficiency is achieved when so

the membrane wall is thicker (Schofield et al., 1990).

A range from 30 to 60 µm is considered as an optimum thickness of MD ((El-Bourawi et

al., 2006). It was proved that the higher thickness, the lower permeate flux is achieved.

Notwithstanding, the very thin thickness of 5 µm of hydrophobic membrane was used in

DCMD configuration (Khayet et al., 2005).

2.12.2 Effects of membrane porosity

The porosity of membrane is a very important membrane parameter which affects directly

on permeate flux. To get higher membrane evaporation surface area, the porosity must be

higher value. For all MD configurations, this value is in the range of 30-85% generally.

2.12.3 Effects of pore size

It is proved that once the membrane pore size increases, the permeate flux also increase

together (El-Bourawi et al., 2006), but once pore size increasing, the ability of pore wetting

becomes greater. In general, membrane that has pore sizes in a range of 100 nm to 1 µm

is used in MD module. Considering on mass transfer aspect, Knudsen diffusion is

responded for the membrane which has pore size very small, and for larger size, it is

Knudsen-viscous in combination (Khayet and Matsuura, 2011). This determination is

based on the comparison between membrane pore size and the mean free path of vapor

molecular. Depending on the different feed solutions (with different characteristics), the

optimum pore size is determined respectively.

2.12.4 Effects of pore tortuosity

Membrane tortuosity factor is determined as the ratio between average length of the pores

and the thickness of membrane. In general, to calculate the permeate flux, the value of

tortuosity factor of 2 is usually assumed (Lawson, Lloyd, 1997). However, the value of 3.9

is also used as the highest value in the consideration on all MD studies which have been

used to implement (El-Bourawi et al., 2006). In fact, a pore does not cut straight throughout

the membrane wall, so the tortuosity is usually higher than 1. A volatile molecule must

transport across a long tortuous pore way, this problem can be approached to a lower

permeate flux.

2.12.5 Effects of the geometric configuration of membrane module

There are 4 type of membrane module geometries used in most of MD studies: Flat sheet,

frame, spiral wound, and tubular (capillary membranes). The requirements of a typical

membrane module are: high feel solution flow rate, high permeability, high turbulent flow,

less pressure lost within module. Another critical aspect is the distribution of feed aqueous

cover the feed side of membrane which mainly effect on mass transfer effective. It can be

said that there is a mal-distribution of feed liquid flow along the membrane surface of

hollow fiber module. This phenomenon is occurred because of the non-uniformity of the

inner diameter at the lumen side and at the shell side of fiber packing.

35

2.12.6 Optimization of membrane distillation process

Experimental designing and operational conditions are two key factors for studying of

optimizing membrane distillation process. There are some validated models which are used

to analysis the variance and predict the permeate flux. By using Monte Carlo method,

Khayet (2012) found out the optimum operational parameters for flat sheet sweeping gas

membrane distillation: feed and sweeping gas temperature of 71.6 °C and 17.3 °C, feed

velocity of 0.16 m/s (flow rate of 165 L/h) and a gas fow rate of 36 L/min. The permeate

flux achieved under these conditions was 2.789×10-3 kg/m2.s. The flat sheet membrane has

ability to reject 99.94% and 99.48% when testing with 30 g/L NaCl solution and seawater

respectively.

2.13 Advantages and Limitations of Membrane Distillation Technology

2.13.1 Advantages of membrane distillation technology

In comparison with conventional membrane separation processes, MD has some significant

ascendants such as: eliminate 100 % of ions, colloids and all non-volatile materials;

operating with lower pressure; other conventional distillation process need to operate with

higher temperature than MD; less impact of chemical from process solutions on membrane

operating process; vapor spaces will be scaled down; minimum external energy source and

land requirement; less membrane fouling condition.

2.13.1.1 Low energy requirement

It was identified that the cost for production process is very sensitive to the temperature of

feed solution. (Criscuoli et al., 2008). The heating period is the main portion of energy

consumption of whole process. In this technology, the feed temperature is lower than

boiling temperature, and lower than other conventional distillation process. So, the energy

requirement is very low.

By installing a heat exchanger in the MD module, the system can be optimized by heat

recovery and les energy requirement (Zuo et al., 2011). Different from conventional

process (MF, UF, RO...), to operate the process, MD does not require a pressured feed flow,

which has a very high energy consumption. Because this process is operated by non-

isothermal driven mechanism, so the operating pressure is approximately zero kPa

(Lawson, Lloyd, 1997). Furthermore, MD can be combined with power plant to recovery

heat waste and treat the wastewater at site, then reuse for operational process of this

industry. The treated water has low temperature and high quality. Heating the temperature

of feed liquid solution higher than its boiling temperature is unessential. The temperature

of feed water is particularly ranged from 60 to 90oC (Khayet, 2011), but 20oC is even

confirmed by El-Bourawi (2006).

On another hand, the cost of MD production process can be reduced by using renewable

energy. The most interesting energy source for MD operating is solar and wind energy, but

the permeate flux is mainly affected by solar intensity (Susanto, 2011).

2.13.1.2 High selectivity and efficiency

36

Relied on the vapor-liquid phase equilibrium, approximate 100% of ions, colloids, and

other non-volatile materials are rejected (Khayet, 2011).

2.13.1.3 Long-term of performance

As mentioned in section 2.9, membrane fouling is not the main problem in operating

process. Moreover, membrane module is not easy to be destroyed due to the low feed

temperature and membrane material characteristic. So that, the operating period before

cleaning and life time of MD module can be very long. The size of pores in MD is larger

than other conventional membrane pores, and the material phase which transfer through

those pores is water vapor, so it is not easy to get clogged. Thus, it is less fouling condition.

2.13.2 Drawbacks of membrane distillation technology

Although MD process has many significant advantages, it still survives some drawbacks

which results in limiting the application of MD process.

The fist, membrane distillation process is not able to separate two or more components in

a mixture that both have high vapor pressure (i.e. water and alcohol). Moreover, MD

currently is mostly applied in food industry. It has a very less application in environmental

treatment. Most of articles that have been published were in laboratory scale.

The second, MD process needs an external energy to heat the feed solution which is not

always available on site. Furthermore, this energy accounts for the highest share of the

operational cost.

Last but not least, the reported permeate flux of MD process is lower than other

conventional membrane process.

2.14 Application of Membrane Distillation

Nowadays, MD is mostly applied in 3 fields: trace volatile organic substrate removal from

water, separate ionic and/or non-volatile compound from water, organic compounds

extraction from liquid solutions (El-Bourawi et al., 2006; Lawson, Lloyd, 1997). The two

main purposes of MD process are: production of distilled water, solutions concentrate. The

significant advantage of this technology is the potential to operate at very high

concentration of non-volatile materials in the feed flow. Typical applications of MD are

shown in details in Table. 2.8

In the industrial processing, MD has variable application, such as: food production (juice

and milk processing), water and wastewater reuse, biomedical (removal of water from

blood), alcohol – water separation, desalination of brackish water, concentrate acids

(mostly sulfuric acid and hydrochloric acid). MD is applied successfully in which process

that need a lower temperature to safeguard the quality of product, for example: MD is very

attractive to fruit processing because its operating with lower temperature resulted in better

quality of fruit juices (Vincenza Calabro 1994) and reduced the needed of high temperature

to disinfect biological fluids in medical side (Sakai et al., 1988).

37

Table 2.8 Typical Fields of MD Application (El-Bourawi et al., 2006)

Application area MD Configuration

DCMD AGMD SGMD VMD

Desalination and pure water production from

brackish water x x x x

Nuclear industry (concentration of radioactive

solutions and wastewater treatments; pure water

production)

x

Textile industry (removal of dyes and wastewater

treatment) x x

Chemical industry (concentration of acids,

removal of VÓs from water, separation of

azeotrophic aqueous mixture such as alcohol/water

mixtures and crystallization)

x x x x

Pharmaceytical and biomedical industries

(removal of water from blood and protein

solutions, wastewater treatment)

x

Food industry (concentration of juices and milk

processing) and in areas where high temperature

applications lead to degradation of process fluids

x x x

2.15 Research Gap

The physical properties of triethylene glycol are similar to that of ethylene glycol. Thus, all

ethylene glycol separation process will also be feasible for triethylene glycol recovery.

Table 2.9 Overview of Various TEG Separation Process

Technology Scope TEG concentration

achieved (%) Reference

Reverse Osmosis (RO) 20 (Jacob et al., 2014)

Nanofiltration (NF) 20 (Jacob et al., 2014)

Pervaporation 99 (Larpkiattaworn et al., 2013)

Distillation 90

Membrane Distillation 78.6 (Yuthawong, 2014)

Remark: - Lab scale/Bench scale, - Full scale

Wastewater with very low TEG concentration (very less than 0.1 %) could be treated by

biological process since it has the biodegradable property. However, it will damage the

38

biological treatment system at higher concentration. Thus, membrane based treatment (NF,

RO, MD), pervaporation and distillation process are proven as the appropriate technologies

to separate TEG from wastewater at high concentration. Table 2.11 presents the membrane

based separations and distillation technology of TEG and their efficiency. RO and NF are

very less potential to concentrate due to the effect of fouling phenomena. When the TEG

concentration increases more than 80 %, pervaporation and distillation process are the two

suitable technologies. Pervaporation process can recover TEG by using partial vapor

pressure and less energy consumption which are similar principle with VMD configuration

of MD process. Distillation process has a very high efficiency (90% TEG concentration

can be achieved), however it requires high energy.

Membrane distillation process has a high potential on separating TEG from wastewater at

the concentration from 0-70 %. Once the volume of TEG wastewater reduces, the total

treatment cost also decreases consequently. However, there is very less study on hollow

fiber SGMD configuration at both bench scale and pilot scale. Up to December 2010, there

is only 4.5 % of MD papers deal with SGMD configuration (Khayet et al., 2012). Almost

studies of other authors focused on flat sheet fabrication at bench scale and conducted with

synthetic wastewater, not on real wastewater. Besides, energy consumption is the very

important aspect that has not been evaluated sufficiently. Thus, the overall effects of

operating variables and energy utilizations should be evaluated holistically.

Some highlight points which are critically concluded from literature

1. TEG is now using as a solvent for the dehydration of moisture of natural gas

processing.

2. Concentration of TEG in wastewater is various from 5 to 20 % and the treatment

cost is very high (4,500 baht/m3).

3. TEG has low risk potential to environment and human health since it can be

biodegradable and less volatility. However, TEG wastewater has some

environmental toxicology problems.

4. Membrane distillation has ability to concentrate TEG up to 78.6%. Nevertheless,

its efficiency and performance can be improved by optimizing operational process.

Economic feasibility of SGMD process has not been evaluated sufficiently.

39

Chapter 3

Methodology

This chapter describes in detail about the materials and methods which were used for the

investigational purpose of the study. This research was conducted in ambient laboratory,

Asian Institute of Technology. The details of specific experiments and all conditions of this

study are discussed in the following sections

3.1 Methodology Overview

There were two phases of this study, namely: (1) bench scale unit and (2) pilot scale hollow

fiber membrane distillation unit study. Both of these units were operated with sweep gas

membrane distillation (SGMD) configuration. The details of this study is shown in Figure

3.1

Firstly, two hollow fiber membrane distillation modules were verified by pure water feed

solution before starting the experiments. The verification process was aimed to measure

the coefficient and membrane resistance. The verification experiments were conducted to

confirm that there were only volatile molecules in permeate flux. Non-volatile compounds

were desired to be separated out at the feed side of the membrane.

Secondly, the membrane modules were run with distilled water. The variables in this step

included: feed temperature, feed flow rate, and sweeping gas flow rate (for the sweeping

gas flow rate, it was kept increasing as high as possible, until the air bubble appeared at the

feed side of membrane or the permeate flux reached highest value). The best-performance

condition was chosen to conduct the next experiment.

In the third step, synthetic wastewater was used as feed solution to run the membrane

system. Synthetic wastewater was a mixture of pure TEG and distilled water. It was

prepared in the ambient laboratory and was used to evaluate the performance of MD

process on concentrating TEG without any interference from other impurities.

Lastly, the membrane systems were operated with real wastewater which was provided by

PTT. In real wastewater, the concentration of TEG was various from 5-20 %. Moreover,

real wastewater was expected to contain other impurities such as benzene, toluene,

ethylbenzene, xylene (BTEX), iron, and other organic compounds. The purpose of running

membrane system with real wastewater was to identify the performance of hollow fiber

membrane distillation in the condition which was close to the real condition at gas

separation plant.

Energy consumption and membrane fouling phenomena were evaluated. The optimum

condition was chosen based on the balance among four factors: permeate flux, final TEG

concentration, fouling, and energy consumption.

The results from this study were used to design a real MD plant, of 1 m3/day capacity.

40

Pre-t

rea

tmen

t

un

it

Real TEG wastewater (5-20%) Pre - treatment unit

(MF + UF system)

Pre-treated TEG wastewater

Water analysis

- Turbidity - Oil and grease

- BTEX - TEG concentration

Pil

ot

sca

le u

nit

Synthetic TEG wastewater

(Based on the best operational

condition from bench scale study)

Analyzing parameters

for all experiments

Flux

Energy consumption

Rejection

Real wastewater

(5-20%)

(Based on the optimum

operational condition of

synthetic TEG

wastewater test)

Hollow fiber (HF) 0.45 µm

System verification

Optimization

(1) TEG concentration (2) Operation of MD

system

Ben

ch

sca

le u

nit

Hollow fiber (HF) 0.45 µm

Distilled water

Synthetic wastewater

TEG concentration (%):

5, 10, 20, 30, 40, 50, 60

Feed temperature (oC):

50, 60, 70

Feed flow rate (L/min):

1.8, 2, 2.4

Sweep gas flow rate

Best

performance

condition

Analyzing parameters for

all experiments

Fouling

Energy consumption

Rejection

Optimization (1) Permeate flux (2) TEG concentration

Real wastewater (5-20%)

(Based on the optimum

operational condition of pure

water and synthetic TEG

wastewater test)

System Verification

Flow rate

Temperature

Rejection

Flow rate

Temperature

Rejection

Figure 3.1 Experimental study plan

41

3.2 Experimental Materials

The main materials used in this research were membrane module and aqueous solutions.

The details are summarized in Figure 3.2.

Figure 3.2 Experimental materials using in this research

3.2.1 Membranes

3.2.1.1 Microfiltration and ultrafiltration membrane

In pretreatment unit, hydrophilic membranes were used. The properties of both MF and UF

membrane used in pre-treatment unit are shown in Table 3.1

Table 3.1 MF and UF Membrane Properties in Pre-treatment Unit

Membrane properties

Details MF UF

Company Mazuma SMMET

Model number Ceramic OBE cartridge UFH-PST-90 (Standard

4040)

Pore size (µm) 0.3 50-60 kD

Membrane area (m2) 11.4 0.47

pH range 5.5-9.5 2-13

Module diameter (mm) 1,400 30

Module length (mm) 5,100 350

Temperature range (oC) 5-38 5-45

*Remark: SMMET- Shanghai Megavision Membrane Engineering and Technology

Research materials

Types of Membranes

- Hollow fiber membrane

distillation (Hydrophobic

membrane)

+ Bench scale: 0.255 m2

+ Pilot scale: 2.0 m2

- MF and UF membrane (pre-

Types of Solutions

- Deionized water

- Synthetic TEG

- Salt solution

- Real wastewater

- Cleaning solutions

42

3.2.1.2 Membrane distillation

The characteristics of membrane used in bench scale and pilot scale unit are shown in Table

3.2. The fabrication of hollow fiber membrane aims to achieve high packing density

(m2/m3). It was difficult to clean the membrane module in this fabrication. Thus, the

pretreatment system was required to work well generating very less solid particles in the

effluent source. In this study, feed solution was fed at outer surface of fiber while sweeping

gas would go into the inner side of fiber (outside-in configuration). The hydrophobic nature

of this membrane avoided liquid to enter the membrane pore.

Table 3.2 Hollow Fiber Membrane Distillation Specification

Descriptions Characteristics

Bench scale unit Pilot scale unit

Company Name Sumitomo Electric Industries, Ltd.

Membrane Name TB-21-02 PM-X215

Type No. 130529-1 13001-2

Module Configuration Hollow Fiber

Membrane Material Polytetrafluoethylene (PTFE)

Type of Membrane Hydrophobic microporous

Contact Angle 112o

Nominal Pore Size (μm) 0.45

Outside Diameter (mm) 2.03

Inside Diameter (mm) 1.07

Total Length (mm) 500 1105

Effective Length (mm) 400 1008

Thickness (μm) 480

Number of fibers 100 306

Membrane Effective Area (m2) 0.255 2

Operating Temperature Range (oC) -100 to 260

pH Range 0-14

Hollow fiber membrane distillation module

The specification of hollow fiber membrane module ws summarized in Table 3.3.

3.2.2 Experimental solutions

3.2.2.1 Deionized water

43

Deionized water (DI) used in this study was produced from the deionization system in

Ambient Lab, Environmental Engineering and Management program, AIT. The electrically

conducting ions are removed from water by using an ion exchange unit.

Table 3.3 Membrane Distillation Module Specification

Description Characteristics

Bench scale unit Pilot scale unit

Type of Membrane Module Hollow fiber

Module Configuration SGMD

Frame Material Polysulfone

Driving Force Thermal driven

Inner Space (cm3) 763.72 7,300

Pipe Diameter (mm) 6 25.4 (feed)/6

(permeate)

Dimension of Module (cm)

Diameter

Length

4.8

40.5

9.6

110.5

Operating Temperature Range (oC) -60 to 90 -100 to 149

pH Range 2-13

3.2.2.2 Salt solution

The salt solution has concentration of 1 %. This chemical were represented the high

concentration of non-volatile substrate in real wastewater source. The salt used in this study

was Sodium Chloride (NaCl) with the purity of 99%.

3.2.2.3 Synthetic TEG

Stock TEG 90% by volume collecting from TEG recovery unit was provided by PTT. It

was stored in 200L tank in the cold storage (5oC). The synthetic TEG wastewater was

prepared by mixing pure TEG and DI water using Equation 3.1.

M1V1 = M2V2 (3.1)

Where Mi, Vi are concentration of TEG and volume of mixed liquid respectively

3.2.2.4 Real wastewater

The real wastewater was collected at gas separation plant and was provided by PTT. It was

expected to have some impurities such as suspended solids, oil and greases, organic

compounds, BTEX (benzene, toluene, ethylbenzene, xylene) and iron. After receiving

44

wastewater from PTT, it was stored in a composite container and was pre-treated before

operating with MD system.

3.2.2.5 Cleaning chemicals

Chemical agents which were used to clean the hydrophobic membrane were a mixture of

oxalic and citric acid. This cleaning solution had been proven as an efficient solution to

clean the hydrophobic membrane distillation (Guillen-Burrieza et al., 2014). The specific

information of the cleaning chemicals are described in Table 3.4.

Table 3.4 Cleaning Chemicals Used in this Study

Generic

Name

Chemical

Formula

MW

(g/mol) CAS No.

Purity

(%) Supplier

Citric Acid C6H8O7 192.12 77-92-9 99.8

U & V Holding Co., Ltd Oxalic

Acid C2H2O4 90.03 144-62-7 99.8

3.3 Experimental Methods

3.3.1 Experimental set-up

Pre-treatment unit

The pre-treatment unit was a combined membrane system, hollow fiber ultrafiltration

membrane (UF) followed by cartridge microfiltration membrane (MF). This unit played

as a physical filtration barrier, to remove suspended solids (by MF), oil and grease (by UF).

By removing of such impurities, pre-treatment unit helped to minimize the particle fouling

phenomena and increase the separating efficiency of MD module. The schematic diagram

of pre-treatment unit is clearly described in Figure 3.3.

In the raw wastewater tank, an air diffuser was installed. The purpose of adding air diffuser

was to remove BTEX and transfer iron ion to ferric hydroxide (Fe(OH)3). The air-to-water

ratio was based on the concentration of iron in wastewater. Beside, viscosity of raw

wastewater also reduced slightly after passing through pretreatment unit. It would lead to

reduce fouling condition.

The parameters which were used to evaluate pre-treatment unit’s performance are:

turbidity, oil and grease, BTEX, TEG concentration, and fouling

Bench scale system set-up

Crossing flow mode was used as the mode operation of this study. Feed aqueous solution

contacted directly and was circulated in cross flow mode over the feed surface of

membrane. Sweep gas contacted the membrane surface and crossed the permeate side. Heat

and mass transfer would cross the hydrophobic porous structure of membrane. It led to

reduce effect of boundary layer resistance. Moreover, the effect of concentration and

45

temperature polarization was also reduced (Khayet and Matsuura, 2011). Figure 3.5

describes in detail about this flow mechanism. The schematic diagram of bench scale

hollow fiber MD unit is described in details in Figure 3.6.

Figure 3.3 Schematic diagram of pre-treatment unit

Figure 3.4 Image of pre-treatment system

Raw wastewater tank

D

Intermediate tank Pre-treated tank

UF Module MF Module

P

Air

P P

Level Sensor D Air Diffuser P Pressure Gauge

Pump Valve

46

Figure 3.5 Cross flow mode description in sweep gas membrane distillation

Figure 3.6 Bench scale hollow fiber sweep gas membrane distillation

Figure 3.7 Image of bench scale hollow fiber SGMD system

Membrane

Permeate

Feed

Sweeping gas

Concentrat

e

Membrane Module

Permeate flow

Control box P Pressure Gauge

Pump Valve

Permeate tank

T

T

P T Heater

Feed

tank

T P

Gas Compressor

MD Module

47

In this unit, the feed tank was contacted directly to the electrical heater. Sweeping gas was

provided from gas line of an ambient air compressor by controlling the pressure meter. By

using sweep gas, the effect of temperature polarization at the permeate side of membrane

was reduced. The temperature of inlet and outlet of both feed and permeate flux were

measured by thermocouples.

Pilot scale unit setup

Both Pilot scale and Bench scale used same membrane type but different surface area and

number of fibers. Figure 3.8 shows the schematic diagram of pilot scale unit in detail. In

this unit, feed solution was heated up until the desired temperature by a heat exchanger.

This equipment helped to reduce energy consumption to heat a high volume of feed

solution. The liquid which was used in the heater was water. The heat exchanging

efficiency was maximized when all feed solution contact directly with boiled water.

Sweeping gas used in this scale was same characteristic with bench scale but higher flow

rate.

Figure 3.8 Pilot scale hollow fiber sweep gas membrane distillation unit

In both scale units, membrane distillation process were operated under configuration of

sweeping gas membrane distillation (SGMD). Figure 3.9 shows in details the principle of

SGMD configuration.

Heat exchanger

Heater

Feed tank

Permeate tank

Air compressor

Hollow

fiber

module

P T

P T

T

P T

P

Box controller T Thermocouple P Pressure Gauge

Pump Valve Flow meter Air filter/water trap

48

Figure 3.9 The principle of sweeping gas membrane distillation configuration

3.3.2 Experimental study

3.3.2.1 Bench scale experimental study

System verification

The temperature calibration was conducted to achieve an accuracy of operating temperature

of bench scale unit. Hence, all thermocouples were calibrated by using standard

thermometer before starting operating the experiments

Membrane rejection measurement was conducted to evaluate the performance of MD

system. To ensure that only volatile molecules of desired compounds can pass the

membrane and there would be no wetting phenomenon occurs in the membrane pores.

From Laplace equation (Lawson and Lloyd, 1997), higher operating temperature leads to

lower surface tension (γL). In this study, the highest feed temperature was 70oC, so the

lowest LEP would be achieved. Thus, it was resulted in deterioration of salt rejection and

the risk of membrane wetting would be highest at this temperature. Laplace equation is

shown in Equation 3.2.

𝐿𝐸𝑃 > ∆𝑃𝑖𝑛𝑡𝑒𝑟𝑓𝑎𝑒 = 𝑃𝑙𝑖𝑞𝑢𝑖𝑑 − 𝑃𝑣𝑎𝑝𝑜𝑟 =−2𝐵𝛾𝐿 𝑐𝑜𝑠 𝜃

𝑟𝑚𝑎𝑥 (3.2)

Overall, in membrane rejection test, saline solution 1% was used as feed solution at

temperature of 60oC. By measuring the conductivity of permeate, salt rejection was

evaluated.

Deionized (DI) water was used as a feed solution of MD system in the process of measuring

pure water flux (PWF). The liquid level in feed tank was recorded at every hour. It

decreased continuously in the MD operating process. Amount of decreasing water was

considered as amount of water vapor pass the membrane pores. The value of permeate flux

(J) was described in kg/m2.h . Equation 3.2 shows the formula to calculate the value of

permeate flux.

Warm Feed

Sweeping

Gas

Porous Membrane

49

𝐽 =𝑉𝑗−𝑉𝑖

𝐴×𝑡 (3.3)

Where V𝑖 , Vj, A, t are the volume of solution in feed tank at time i and j, membrane surface

area and time of measuring respectively.

From PWF, membrane coefficient (Bw) and membrane resistance (Rm) was calculated by

using Equation 3.4. Rm was the resistance of a membrane when it was still not yet affected

by other resistances (i.e. boundary layer resistance). Rm of pure feed water was used to

compare with other feed solution to evaluate fouling condition.

𝐽 = 𝐵𝑤 × ∆𝑝𝑊 = 1

𝑅𝑚× ∆𝑝𝑊 (3.4)

Temperatures of feed pure water were 50, 60 and 70oC, feed flow rates were 1.8, 2 and 2.4

L/min. At the permeate side, the sweeping gas flow rate was depended on the pressure, it

could be high as possible if there was no undesirable pressure developed. The operation

period of each batch was 8 hours and data collection was conducted at every hour.

Synthetic TEG wastewater test

Before testing with real TEG wastewater, bench scale hollow fiber SGMD unit was tested

with synthetic wastewater first. From synthetic wastewater study, the optimum TEG

concentration was found out without affecting of any other impurities in feed solution.

There were 7 levels of initial TEG concentration, obtains 5, 10, 20, 30, 40, 50, 60 %.

At each initial TEG concentration, the experiment was run totally 8 hours. The intention

of choosing 8 hours was to evaluate the system’s performance within time of official

working day. The permeate flux was measure at every hour of this 8 hours. For TEG

sample, there were two different samples: initial concentration sample and the sample of

TEG solution after 8 hours of operating.

Real TEG wastewater test

After testing with synthetic wastewater, the bench scale MD unit was test with real

wastewater. There are many other impurities in real wastewater, especially suspended

solids, oil and grease, BTEX. These impurities would increase the opportunity of

membrane fouling. A MF combined with UF system was the appropriate pre-treatment

system for this wastewater. The TEG concentration in real wastewater, in fact, was various

from 5 to 20%, it was uncontrollable. All the measurements of real wastewater testing were

same with synthetic wastewater testing.

3.3.2.2 Pilot scale experimental study

The optimum operating condition was chosen from bench scale hollow fiber SGMD study

(base on permeate flux, TEG concentrating ability, and energy consumption). In the second

phase of this study, the optimum condition from first phase study was applied to operate

the pilot scale hollow fiber SGMD unit. From the experimental results, a full scale SGMD

plant was designed and was accepted by the gas separation plant.

50

System verification

In this phase, only feed temperature at the best- performance condition from bench scale

study was used to operate the pilot scale unit. The feed flow rate, and sweeping gas flow

rate were changed to adapt with a larger scale in comparison with the dimension of bench

scale hollow fiber module and membrane area.

From pure water flux, the coefficient of membrane distillation (Bw) and membrane

resistance (Rm) in pilot scale were calculated. These values present the properties of

membrane in pilot scale and was used for further study on membrane fouling phenomenon.

In the rejection test of pilot scale, due to the reaction between salt (NaCl) and copper (Cu)

in the heat exchanger, TEG solution (10%) was selected. By measuring the TEG

concentration (TEG played as non-volatile compound) at both feed solution and permeate

condense, the rejection ability of membrane was calculated. The evaluation on rejecting of

undesirable compound was a need for every membrane distillation experiment. The pilot

system must be ensured that the pores do not get wetting and only desired compound can

pass the hydrophobic porous membrane. According to Equation 3.2, 70oC was the feed

temperature used in rejection test.

Synthetic TEG wastewater test

The concentration of synthetic TEG wastewater was used in pilot scale unit was the TEG

concentration used in bench scale unit study. Equation 3.1 was used to prepare this

concentration of synthetic TEG wastewater. Permeate flux and energy consumption were

measured at every hour of 8 hours experiment. Membrane fouling phenomena was

evaluated from the value of membrane coefficient in this experiment.

Real TEG wastewater test

Similar with real TEG wastewater test in bench scale unit, pilot scale SGMD was also

tested with real wastewater. The reason for testing with real wastewater was due to the

reality that wastewater contains impurities and uncontrollable concentration. Real

wastewater was stored in a feed tank of 52 L and supplied continuously to the hollow fiber

module for whole 8 hours of experiment. The sample of real wastewater was taken at three

points of timeline: before pre-treatment unit, before operating the pilot scale unit, and every

8 hours of operation. Both real wastewater and Synthetic TEG experiments were test with

two operating method: batch experiment (8 hours per batch) and continuously-fed

experiment (40 hours, equal to 5 working days per week)

3.3.3 Energy consumption evaluation

Energy consumption by MD system was measured by the power meters. At each scenario

of operation, the value of ratio between energy consumption and permeate flow rate would

be considered. The lowest value means highest energy efficiency. The ratio between

energy consumption and permeate flow rate were described in a unit of (kW/kg.h) by using

Equation 3.5.

E =∑ Power consumed (kW)

Permeate flux (kg m2⁄ h)×Membrane area (m2) (3.5)

51

Where E is energy efficiency (kWh/kg)

3.3.4 Membrane fouling

Membrane resistance (Rm) was measured by operating a new membrane with feed solution

as DI water. This was a certain resistance that every membrane owned. When operating

with wastewater (either synthetic or real wastewater), boundary layer was formed on the

membrane surface. Consequently, fouling resistance (Rf) would contribute to total

resistance (Rt). Rf would result in decreasing of permeate flux along with timeline of

operational process. Some deposits could only be removed by chemical cleaning solution.

Besides, the cleaning solutions also help to neutralize the membrane surface before

conducting new experiment.

Time of cleaning period and chemical concentration are shown in Table 3.5. From cleaning

process, each type of fouling resistance were investigated. After operating with wastewater,

both units (bench scale and pilot scale) were operated with DI water. Recoverable fouling

resistance (Rr) was estimated from this step. Then by operating with chemical solutions,

reversible fouling resistance (Rre) was evaluated. Equation 3.6 expresses the relation among

all fouling types. From Equation 3.6, Irreversible fouling resistance (Rir) was calculated.

𝑅𝑡 = 𝑅𝑚 + 𝑅𝑟 + 𝑅𝑟𝑒 + 𝑅𝑖𝑟 (3.6)

3.3.5 Design of full scale SGMD plant

Based on the results of this study, a full scale SGMD system was designed and was

accepted by the Gas Separation Plant. The operational capacity of this system was 1

m3/day. All the costs (investment, operation and maintenance) were estimated based on the

market price at the designing stage. Operational conditions were the optimum conditions

from this study. Real wastewater would be treated and TEG would be concentrated at

lowest-cost as possible. The products of this design included all calculations and Autocad

drawings.

Table 3.5 Membrane Cleaning Procedure

Step Cleaning

solution

Concentration

(mg/L)

Cleaning time

(h) Purpose

1 Water (1) - 0.5 Recoverable fouling

2 Oxalic/Citric 0.1wt. % Oxalic cid

0.8wt. % Citric acid 6 Reversible fouling

3 Water (2) - 0.5 Washing the remain

chemical

3.4 Experimental Analysis

The methods to measure TEG and BTEX concentration in real wastewater using in this

study were provided by PTT.

52

3.4.1 TEG analysis

Firstly, the sample of TEG wastewater wass diluted in methanol solution. After that, the

Gas Chromatograph (GC) method was used to measure TEG concentration. The equipment

to use this method has a name of GC Agilent HP 6890. This equipment was a product of

Agilent Technologies, Inc., USA. The flow rate in Restex Rxi 624 sil MS column was

controlled at 1.2 mL/min. Splitless mode was set up for the Inlet at 250oC. The standard

curve of TEG concentration was in range of 0.1-2.0 % (v/v). The value of R2 (goodness-

of-fit) in this measurement was 0.9993.

3.4.2 BTEX analysis

A SCION Triple Quadrupole (TQ) detector was used to measure BTEX in wastewater

sample. This machine was a product of Brucker Corp., Germany. The technique of this

method was using the headspace sampling. The sample of wastewater was diluted in

distilled water then put in the cylindrical container of Milli-Q before heating up until 80oC.

For supporting the volatilization of BTEX, besides the high temperature, the sample was

shaken at 250 rpm within 30 min. The vapor of BTEX volatile molecules above liquid part

was measured by an injector. The column BR-5 FS that has dimensions of 30 mm x 0.25

mm x 0.25 um was used. The material of this column contains 5% diphenyl and 95%

dimethylpolysiloxane. For gas chromatograph, the highest temperature was 230oC. The

retention time of volatile molecules was one of the important points that needed to be

considered while doing this measurement. The retention time of 2.77, 4.60, 7.19, 7.45, 8.19

were belonged to benzene, toluene, ethylbenze, m,p-xylene, o-xylene respectively. In this

measurement, the value of R2 of standard curve was 0.999.

Figure 3.10 Determination process of fouling resistances

Pure water test (1)

TEG wastewater test

Oxalic/Citric cleaning

Rm

Rt

Rr

Rir + Rre

Clean by DI water

Pure water test (3)

Clean by DI water

Pure water test (2)

53

3.4.3 Other wastewater parameters analysis

Wastewater analysis was conducted to determine the level of pre-treatment system needed

and performance of this unit. From the experimental results, a pre-treatment unit was

designed and set up. The water parameters which were measured are shown in Table 3.6.

TEG has a very high of boiling point and viscosity. During the measurement process of

suspended solid, it was impossible to separate with particle solids by filter paper. TEG

attached particle molecules. At the temperature using in heating sample to measure SS,

TEG could not be evaporated. The experimental value of suspended solid would be

affected. Thus, in this case, turbidity measurement was the most appropriate method to

measure suspended solids parameter.

In the experimental process, synthetic wastewater was also used to run the MD system

before experimenting with real wastewater. By mixing stock TEG (90%) and DI water with

a determined amount, synthetic wastewater was formed at a desired concentration of TEG.

Table 3.6 Analytical Parameters and Methods

Parameter Unit Method Equipment Interference Reference

pH Electrometric pH meter Oily material,

temperature

EPA (2005),

9040 C

Turbidity NTU Nephelometric Turbidimeter Stray light,

air bubble

APHA et al.,

(2005), 2130 D

COD mg/L Titration Chlorine APHA et al.,

(2005), 5220 D

Oil and

grease mg/L Gravimetric

Organic

solvents

APHA et al.,

(2005), 5520 D

Fe 2+ mg/L Spectrometric Spectrometer Base material APHA et al.,

(2005), 3111 B

Volatile

compound

Refer to

section 3.4.2

SCION

triple

Quadrupole

Benzene

Toluene

Ethyl-

benzene

Xylene

Subjected

compound

Refer to

section 3.4.1

GC Agilent

HP 6890

Tri-ethylene

glycol

*Remark: All samples were preserved by keeping in cool-storage at 5oC.

54

Chapter 4

Results and Discussions

In this study, operating parameters for enhancing the performance of membrane distillation

for treating TEG are discussed (in terms of permeate flux and energy consumption). This

chapter presents the results of wastewater analysis, efficiency of pre-treatment system and

optimizing result of MD systems. The hydrophobic hollow fiber membrane (pore size of

0.45 µm, thickness of 480 µm) was studied at two scales: (1) bench scale (0.255 m2) and

(2) pilot scale (2 m2), each scale was investigated with three operating phases. The first

phase focused on salt rejection and pure water flux investigation. In the second phase, the

membrane modules were investigated using synthetic TEG solutions at different

concentrations. Lastly, real wastewater was used to finalize the performance of MD

modules. The optimum operating conditions were selected based on the interrelationship

between permeate flux and energy consumption.

4.1 TEG Wastewater Characterization

On the process of applying membrane distillation technology for concentrating TEG

wastewater from gas separation plants waste stream, the wastewater characterization step

was carried out to identify the necessary level of pretreatment. The pretreated system

removed the undesired impurities to protect and enhanced the performance of MD system.

The analytical results analyzed by ALS Company are expressed in Table 4.1.

Table 4.1 Wastewater Analytical Results

Parameter Unit Result

Sample 1 Sample 2 Sample 3

Wastewater parameter

pH - 4.41 -

COD mg/L 114,000 151,500 94,500

Oil and grease mg/L 15 1 <3

Total suspended solids mg/L 3 11 <5

Iron mg/L 0.1 41 1.2

Volatile compounds

Benzene mg/L - 26.6 -

Toluene mg/L - 14.1 -

Ethylbenzene mg/L - ND -

m,p-xylene mg/L - 7.4 -

o-xylene mg/L - 14.7 -

From the analytical results, some highlight technical issues that could be deduced were the

effects due to oil and greases, BTEX, and iron on membrane distillation process. As pre-

55

treatment, total suspended solids (TSS) and oil and grease (O&G) had to be removed and

be eliminated since they could foul the membrane surface (and/or membrane module). The

concentration of iron as is observed in sample 2 (41 mg/L) was another problem. During

concentrating TEG wastewater, the solubility limit of iron in the concentrated solution

would be reached and iron could potentially precipitated on the membrane surface.

However in sample 1, 2 and 3, high iron concentration was not always observed (0.1 mg/L

in sample 1 and 1.2 mg/L in sample 2). This fluctuation could have occurred due to a

process instability at the methane production site at the GSP. Another observed issue is that

the appearance of volatile compounds in wastewater. Volatile compounds such as benzene,

toluene, ethyl benzene and xylene (BTEX) were found in the wastewater (Table 4.1),

which were also categorised as carcinogen compounds. During operating MD process, such

compounds would pass through the membrane pores and could reduce the hydrophobicity

of the membrane and also cause health effects to the operators. Practically, by providing

aeration into the wastewater prior pre-treatment unit, BTEX could be vaporized due to high

volatility and iron could be precipitated through oxidation reaction. Oil and grease are the

organic species which contains hydrophobic components in the molecule structure. This

such molecules can readily attach onto the membrane surface and lead to process failure in

MD. This could lead to the formation of hydrophilic tracks through the pores. This

approach allows water to (a polar specie) interact with the hydrophilic section of oil and

grease. Thus, it would be a potential risk of liquid penetration into the pores (Padaki et al.,

2015). Overall, it was necessary to install a pre-treatment system prior MD unit. A proper

pre-treated wastewater could lead to an effective performance of MD process for long term

operation.

4.2 Performance of Pre-treatment Unit

As presented in Section 4.1, the pre-treatment system was installed to eliminate suspended

solids, oil and grease predominantly. Such undesired impurities might cause membrane

fouling issues and affect the hydrophobic characteristic. The performance of the process

was determined by comparing the quality parameters such as: raw wastewater (before pre-

treatment unit), effluent from microfiltration membrane (MF), and pre-treated wastewater

(after passing through ultrafiltration – UF). The measured parameters and analytical results

are shown in Table 4.2

Table 4.2 Analytical Results of Pre-treating Samples

Parameters Raw Wastewater After MF After UF

COD (g/L) 151.5 157.9 150.50

Turbidity (NTU) 30.69 4.21 1.27

Suspended solids (mg/L) 11 <5 <5

pH 4.41 4.48 4.46

O&G (mg/L) 15.00 7.00 5.00

56

Figure 4.1 Removal efficiency of pre-treatment system

TEG was represented in wastewater through COD concentration. The analytical results of

COD parameter in three samples (Table 4.2) were almost similar. It could be concluded

that TEG was not separated from wastewater during pretreating process. Microfiltration

(MF) was designed to remove suspended solids (represented as turbidity parameter) while

ultrafiltration responded to eliminate oil and grease. According to Figure 4.1, turbidity

and oil and grease were removed with removal efficiency of 95% and 66 % respectively.

Overall, the targeted impurities (suspended solids, oil and grease) could be greatly removed

after passing through designed pretreatment unit.

4.3 Optimizing Operation of Bench Scale Hollow Fiber SGMD

Bench scale membrane distillation system used 0.255 m2 hollow fiber membrane (0.45 μm)

as described in Section 3.2. The packing density of this membrane was 333.8 m2/m3. During

operating process, the pressure at permeate side (created by sweeping gas) was maintained

to be close to the pressure at feed side (created by feed solution). Regarding to feed flow

rate, if it excess the liquid entry pressure (LEP), liquid would pass through the membrane

pore causing pore wetting phenomena.

4.3.1 MD process verification

The purpose of doing pure water test was to ensure the system operated according to the

principles of MD process. MD coefficient and MD resistance were evaluated from this

experiment. In MD process, there was only vapour of volatile material crossing throughout

the hydrophobic micro-porous membrane. Thus, pure water flux and salt rejection were the

two main experiments to verify this step.

4.3.1.1 Rejection test

Before operating the bench scale hollow fiber SGMD with feed solutions (i.e pure water,

synthetic TEG, real wastewater), the rejection test (with 1% salt concentration) was

conducted to ensure that the membrane modules were working well with membrane

distillation mechanism. In principle, the hydrophobic membrane distillation only allowed

0

20

40

60

80

100

Turbidity (NTU) O&G (mg/L)

Rem

oval

Eff

icie

ncy

(%

) After MF After UF

57

volatile molecules to pass through the membrane pores. Therefore, the conductivity of

condensed permeate should not increase. The rejection ability of hollow SGMD was

investigated at feed flow rate of 2.4 L/min, feed temperature at 70oC and sweeping gas

velocity of 3.1 and 5.3 m/s. Feed temperature of 70oC would be the highest temperature

used to operate the membrane module. The higher feed temperature, the lower liquid entry

pressure membrane has. Consequently, non-volatile compound had a higher chance to pass

through the membrane pores. Also increasing the sweeping gas velocity at permeate side

of hollow fiber membrane, the strong sweeping gas had a greater potential of sucking the

liquid from the feed side.

The rejection was calculated using the following equation:

𝑅(%) = (𝐶𝑓−𝐶𝑝

𝐶𝑓) × 100 (4.1)

The experimental results of rejection test are expressed in Figure 4.2 and 4.3.

Figure 4.2 Rejection result of hollow fiber SGMD at gas velocity of 3.1 m/s

As it can be observed from Figure 4.2 and 4.3, the hydrophobic membrane distillation was

able to reject 99.99 % of salt. The conductivity of condensed permeate flow was affected

by the impurities from sweeping gas pipe lines. The results show that there was no non-

volatile compounds penetrate throughout the membrane during operating process. Thus,

the membrane module performed well with MD principle. The reduction of permeate flux

during rejecting process was due to the increase in salt concentration.

0

20

40

60

80

100

1.4

1.6

1.8

2.0

2.2

1 2 3 4R

ejec

tion (

%)

Flu

x (

kg/m

2.h

)

Time (h)

Flux Rejection

58

Figure 4.3 Rejection result of hollow fiber SGMD at gas velocity of 5.3 m/s

4.3.1.2 Pure water test

Hollow fiber membrane module was investigated with DI water. The experimental variables

in bench scale study were feed flow rate at 1.8, 2, 2.4 L/min, feed temperature at 50, 60,

70oC, sweeping gas velocity at 3.1, 3.6, 4.2, 4.7, 5.3 m/s. The driving force in MD process

is the difference in vapour pressure between feed side (liquid phase) and permeate side (gas

phase) of the fiber. Increasing feed temperature, the vapour pressure of water also increases.

Vapour pressure strongly affected by the feed temperature. At the same feed flow rate, same

sweeping gas velocity but lower feed temperature, the permeate flux was lower

significantly. When increasing the sweeping gas velocity, permeate flux also corresponding

increases too. However, as the vapour pressure increases at permeate side this leads to a

decrease in the driving force of MD process. When the sweeping gas velocity exceed a

certain value, the pressure at permeate side is higher than feed side. At this situation, some

air bubble started to appear at the feed side of the membrane.

Effect of feed inlet flow rate

Figure 4.4 presents the effects of various feed inlet flow rates on permeate flux at different

sweeping gas inlet velocities and feed temperature of 70oC. At the same operating conditions

(gas velocity, feed flow rate, feed temperature), the achieved permeate flux were almost

similar. The small difference of permeate fluxes at the sweeping gas velocity were due to

the slightly reduction of boundary layer resistance at the feed membrane surface.

0

20

40

60

80

100

2.0

2.3

2.6

2.9

3.2

1 2 3 4

Rej

ecti

on

(%

)

Flu

x (

kg/m

2.h

)

Time (h)

Flux Rejection

59

Figure 4.4 Permeate flux at different feed inlet flow rate and sweeping gas velocity

at feed temperature of 70oC

Figure 4.5 Permeate flux at different feed inlet flow rate and sweeping gas velocity

at feed flow rate of 2.4 L/min

Practically, it has been proved that the permeate flux in SGMD configuration is not really

affected by the feed flow rate. It could be considered as negligible (Khayet and Matsuura,

2011). This is due to the fact that the sweeping gas flow rate at permeate side governs the

temperature polarization effect during MD process (Khayet et al., 2002). The negligible

effect of feed flow rate was also discussed by other authors (Ding et al., 2006; Zhao et al.,

1.0

1.4

1.8

2.2

2.6

3.0

3.4

3.1 3.6 4.2 4.7 5.3

Per

mea

te f

lux

(kg/m

2.h

)

Sweeping gas velocity (m/s)

1.6 L/min 2 L/min 2.4 L/min

0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

3.1 3.6 4.2 4.7 5.3

Per

mea

te f

lux

(kg/m

2.h

)


50°C 60°C 70°C

60

2014). The mass transfer coefficient was only slightly increased when using ammonia

aqueous as feed solutions for SGMD process. Moreover, Lee and Hong (2001) who studied

the effect of feed flow rate on PTFE hollow fiber SGMD, used IPA feed solutions. Once the

feed flow rate increased, the overall permeate flux slightly enhanced, but not that much. The

reason was due to the decrease of feed boundary layer resistance (higher turbulence flow).

However, more attention to potential risk of pore wetting should be considered with the

increase in liquid hydrostatic pressure (Khayet and Matsuura, 2011).

Effect of sweeping gas velocity

Figure 4.4 and 4.5 presents the variation of permeate flux during the optimization process

of bench scale SGMD system (varied feed temperature and sweeping gas flow rate while

feed flow rate was kept constant at 2.4 L/min). As discussed in Section 4.3, the effect of

feed flow rate was negligible for SGMD configuration.

With the same area section of inside space of hollow fiber, sweeping gas flow rate could be

understood by the term of gas velocity. With an increase in gas velocity, it was noticed that

the hydrostatic pressure which was created by sweeping gas was to be kept lower than than

the corresponding value in the feed solution. Otherwise, the volatile molecules could not

pass through the pores. To avoid pore wetting phenomena, the liquid entry pressure (LEP)

should be higher than the transmembrane hydrostatic pressure. As a result, it was necessary

to identify the optimum gas velocity to maximize the permeate flux in SGMD process. With

any increase in gas velocity would lead to decrease in permeate flux. Increasing gas velocity

resulted in decreasing temperature polarization effect and increasing heat transfer

coefficient.

As it can be observed from both Figure 4.4 and 4.5, the sweeping gas velocity of 4.7 m/s

(equal to the flow rate of 22.5 L/min) was the optimum value. Combining with other

operational parameters, the highest permeate flux achieved was 3.14 kg/m2.h. In this study,

sweeping gas used in SGMD process was the atmospheric air which was provided by the

air compressor. Therefore, there was no control in the sweeping gas inlet temperature. The

usual gas inlet temperature was in the range from 29 to 35oC. It was varied during day- time

and nigh-time and was affected by the season.

4.3.1.3 Energy consumption in pure water test

Figure 4.6 performs the average energy consumption at different scenarios for pure water.

In bench scale MD module, the energy was consumed by 3 main components: feed pump,

heater and gas compressor. The energy consumed by control box was very small and can be

considered negligible. To balance energy consumption and permeate flux achieved, the ratio

of energy consumption/flux was evaluated and the data is summarized in Table 4.3. The

optimum operating condition with feed temperature of 70oC, sweeping gas velocity of 4.7

m/s, the energy consumption/Qp ratio was observed to be the lowest (1.09 kW/kg.h-1).

61

Figure 4.6 Energy consumption/Qp at feed flow rate of 2.4 L/min

In literature MD configuration like DCMD and VMD achieved this ratio at 3.55 kW/kg.h-1

and 1.1 kW/kg.h-1 respectively (Criscuoli et al., 2008). Higher temperature, the heat tank

consumed higher energy. However, the permeate flux was also very high. Then, the ratio of

energy consumption/Qp reduced. In contrast, at lower temperature, the energy consumption

per hours was lower and the energy ratio (kW/kg.h) was higher. Table 4.3 shows the

experimental value of energy consumption analysis.

Table 4.3 The Ratio of Energy Consumption/Qp at Different Scenarios of Pure Water

Test

Sweeping gas

velocity

(m/s)

Feed

temperature

(oC)

Permeate flux

(kg/m2h)

Energy Consumption/Qp

(kWh/kg)

3.1

50 0.52 2.70

60 1.31 1.39

70 1.83 1.35

3.6

50 0.78 1.99

60 1.57 1.28

70 2.09 1.32

4.2

50 1.05 1.58

60 1.83 1.34

70 2.61 1.07

4.7 50 1.05 1.85

60 1.92 1.37

1.1

0.6

1.0

1.4

1.8

2.2

2.6

3.1 3.6 4.2 4.7 5.3

Ener

gy C

onsu

mpti

on (

kW

/kg

)


50°C 60°C 70°C

62

70 3.14 1.09

5.3

50 1.05 2.14

60 1.57 1.64

70 2.7 1.30

Overall, the best condition in terms of balancing between energy consumption/Qp and

permeate flux were achieved at feed flow rate of 2.4 (L/min), feed temperature of 70oC, and

sweeping gas velocity of 4.7 m/s. At this condition, pure water flux achieved was 3.14

kg/m2.h, the ratio of energy consumption/Qp was 1.09 kW/kg. A comparison between pure

water flux obtained in this research and observed in literature is presented in Table 4.4.

4.3.1.4 Membrane coefficient and resistance

By using the result of pure water test, membrane coefficient (Bw) and membrane resistance

(Alkhudhiri et al., 2012) would be evaluated. Equation 2.1 shows the relationship between

permeate flux and membrane coefficient. Rw can be calculated by using Equation 4.1.

𝐑𝐰 =𝟏

𝐁𝐰 (4.1)

All operating temperatures used to calculate vapour pressure are temperature at membrane

surface (Tmf and Tmp) (Equation 4.3 and 4.4). This is due to the vaporization phenomena

taking place at the entrance of membrane pore. Thus, it is necessary to calculate Tmf and Tmp

to evaluate Bw more exactly. The procedure to calculate Tmf and Tmp is expressed as below

(Martı́nez-Dı́ez and Vázquez-González, 1999):

1) Calculate the value of Reynolds number (Re), Prandtl number (Pr) and Nusselt

number (Larpkiattaworn, 2013) at both feed side and permeate side. From those

values, heat transfer coefficient of boundary layer at feed surface (hf) and permeate

surface (hp) are evaluated.

2) Assume Tmf and Tmp then calculate the vapour heat transfer (hv) by Equation 4.2.

𝐡𝐯 =𝐉 × 𝛗

𝐓𝐦𝐟 − 𝐓𝐦𝐩

(4.2)

Where φ is the latent heat of water vapour evaporation (2270 kJ/kg).

3) Use the calculated hv to evaluate Tmf and Tmp by two Equations below:

63

Table 4.4 PWF Comparison with Other Studies on SGMD

Membrane type Gas velocity

(m/s)

Tf

(oC)

PWF

(kg/m2.h) Reference

Hollow fiber

(PTFE – 0.45µm)

3.1

50 0.52

This study

60 1.31

70 1.83

3.6

50 0.78

60 1.57

70 2.09

4.2

50 1.05

60 1.83

70 2.61

4.7

50 1.05

60 1.92

70 3.14

5.3

50 1.05

60 1.57

70 2.7

Hollow fiber (PTFE) 8.7 45 1.95 Lee and Hong,

2001

Flat sheet

(TF-200)

1.0 65 5.4

Khayet et al.,

2000

1.5 65 7.2

2 65 9.0

Flat sheet

(TF-450)

1.0 65 6.8

1.5 65 8.6

2 65 10.0

𝐓𝐦𝐟 = 𝐓𝐛𝐟 − (𝐓𝐛𝐟 − 𝐓𝐛𝐩) ×𝟏 𝐡𝐟⁄

𝟏 (𝐡𝐦 + 𝐡𝐯) + 𝟏 𝐡𝐟⁄ + 𝟏 𝐡𝐩⁄⁄ (4.3)

𝐓𝐦𝐩 = 𝐓𝐛𝐩 + (𝐓𝐛𝐟 − 𝐓𝐛𝐩) ×𝟏 𝐡𝐩⁄

𝟏 (𝐡𝐦 + 𝐡𝐯) + 𝟏 𝐡𝐟⁄ + 𝟏 𝐡𝐩⁄⁄ (4.4)

1) Repeat the procedure until the difference between the assumed calculated values

does not exceed 0.1%.

64

Tmf and Tmp are different from Tbf and Tbp due to the effect of flow condition and heat

transfer efficiency. This phenomena is named as temperature polarization. At Tbf = 69 and

Tbp =48 oC, Tmf and Tmp are 66.60 and 64.53oC respectively. The permeate temperate is

higher in the bulk than at the membrane surface is due to the fact heat transfer was occurring

from the feed side and also that air gains heat faster than liquids.

Table 4.5 Membrane Surface Temperature and Temperature Polarization Coefficient

Feed Permeate TPC

Tbf (oC) Tmf (oC) Tbp (oC) Tmp (oC)

69 66.6 48 64.5

0.1 59 57.2 43 55.6

49 47.7 38 46.6

*Remark: Tbf and Tbp are bulk feed and bulk permeate temperature.

Table 4.6 Experimental Membrane Distillation Coefficient and Membrane Resistance

Membrane Configuration Mass Transfer

Mechanism

Bw

(10-8 s/m)

Rw

(107 m/s)

Hollow fiber

PTFE

(0.45 µm)

SGMD

(This study) Knudsen and

Molecular

diffusion

combined

3.47 2.9

DCMD

(Yuthawong,

2014)

4.2 23.6

As presented in Table 4.6 membrane coefficient of hollow fiber PTFE (0.45µm) operating

with SGMD configuration was 3.47 x 10-8 s/m and was higher than while operating as

DCMD configuration. Yuthawong (2014) achieved Bw at 4.2 x 10-8 s/m. Implying that

hollow fiber PTFE MD had a higher permeability when operating with SGMD configuration

(lower membrane resistance).

4.3.2 Synthetic TEG

4.3.2.1 Batch experiments

Before investigating MD system with real TEG wastewater, it was important to test with

synthetic TEG wastewater. This experiment was aimed at investigating MD performance

with only TEG, without BTEX and any inorganic foulant. In synthetic wastewater test, the

effect of boundary layer resistance was also evaluated. Boundary layer was formed by the

appearance of TEG in feed solution. However, it theoretically creates less effect in SGMD

configuration (Khayet and Matsuura, 2011).

65

The synthetic TEG wastewater was prepared by mixing stock TEG 90% with DI water with

a determined concentration. The bench scale SGMD module was operated at the optimum

condition which is selected from pure water test: Feed temperature of 70oC, feed flow rate

of 2.4 L/min and sweeping gas velocity of 4.7 m/s, changing the initial TEG concentration.

The actual and theoretical TEG concentrations are shown in Table 4.7.

As presented Figure 4.7, when the hollow fiber MD was tested at 10% initial TEG

concentration, the permeate flux was stable at 2.61 kg/m2.h. After the TEG concentration in

feed solution reached 25%, the permeate flux reduced to 2.1 kg/m2.h. During concentrating

TEG from 25 to 45 %, the permeate flux was approximate 2.1 kg/m2.h.

Table 4.7 TEG Concentration in Synthetic TEG Test

No

Theoretical TEG concentration

(%)

Analysis TEG concentration

(%)

Initial Final Initial Final

1 10 34.9 10.59 44.59

2 30 70.4 31.28 60.35

3 60 87.5 61.08 90.52

Figure 4.7 Permeate flux and TEG concentration at synthetic TEG 10% initial

concentration

5

15

25

35

45

55

1.5

2.0

2.5

3.0

0 2 4 6 8 10 12 14

TE

G C

once

ntr

atio

n (

%)

Per

mea

te f

lux (

kg/m

2.h

)

Time (h)

Permeate flux TEG Concentration

66


concentration

It could be observed from Figure 4.8 that the average permeate flux was 2.1 kg/m2.h during

the first 9 hours. In this period, TEG concentration increased from 30 to 55 %. After that,

the permeate flux reduced to 1.5 kg/m2.h since TEG concentration higher than 55 % due to

the expanding of boundary layer and feed viscosity.

The driving force of MD process is affected by the vapour pressure at feed side and permeate

side. The sweeping gas inlet temperature at 6:00 hrs, 12:00 hrs and 18:00 hrs are 28, 34,

30oC respectively. At higher temperature, the humidity would be higher. As a result, the

vapour pressure at permeate side was observed to be higher. It caused a reduction in the

driving force of membrane distillation. During the concentration process, TEG

concentration increases leading to increasing value of thermal conductivity (higher

viscosity). The surface temperature at feed side does too. Consequently, the vapour pressure

at permeate side increases which enables to increase the driving force of this process. When

the TEG concentration continuously increased from 55 to 75%, permeate flux was about 1

kg/m2.h. At higher TEG concentration (>75%), permeate flux was very low (less than 0.5

kg/m2.h). At this situation, MD was affected by the feed viscosity. In SGMD configuration,

the effect of concentration polarization very small in comparison with that of temperature

polarization (Khayet and Matsuura, 2011; Lawson and Lloyd, 1997).

25

35

45

55

65

75

1.2

1.6

2.0

2.4

2.8

0 2 4 6 8 10 12 14

TE

G C

on

cen

trat

ion

(%

)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


67


concentration

4.3.2.2 Continuously-fed synthetic TEG experiment

Synthetic TEG (10%) was prepared by mixing stock TEG (90%) with DI water. It was used

to fill up the feed tank every 8 hours in this experiment. The TEG concentration reduced

dramatically at the moment of feeding by synthetic TEG 10%. However, its concentration

increased in whole process (from 10% to 53%) because of the continuous feeding progress.

All experimental results of this experiment is shown in Figure 4.10.

Figure 4.10 Experimental result of continuously-fed synthetic TEG investigation

55

65

75

85

95

0.0

0.4

0.8

1.2

1.6

0 5 10 15 20 25

TE

G C

on

cen

trat

ion

(%

)

Per

mea

te f

lux

(kg/m

2. h

)

Time (h)


5

15

25

35

45

55

1.5

2.0

2.5

3.0

0 10 20 30 40 50

TE

G c

on

cen

trat

ion(%

)

Per

mea

te f

lux (

kg/m

2.h

)

Time (h)

Permeate flux

TEG concentration

68

This experiment was run 40 hours continuously, feeding every each 8h without draining

out any concentrated feed solution. The permeate flux was stable at 2.6 kg/m2.h when the

TEG concentration was less than 25%. When the TEG concentration was in the range from

25% to 35%, permeate flux reduced from 2.6 to 2.4 kg/m2.h. At higher TEG concentration

(>35%), the flux varied around 2 kg/m2.h as seen in Figure 4.10. The explanation for

reducing of the permeate flux is due to the increasing of boundary layer resistance when

TEG concentration increases (higher viscosity). All experimental data is shown in

Appendix E

4.3.2.3 Boundary layer resistances

It was observed that at the same operating condition of feed flow rate, feed temperature and

sweeping gas flow rate, the PWF was 3.14 kg/m2.h with the membrane resistance being 266

x 105 m/s. Table 4.8 presents the resistance of membrane and boundary layer. Increasing

TEG concentration, the effect of boundary layer resistance contributed more towards

reducing the permeate flux. The experimental boundary layer resistance are also presented

in Table 4.8.

The boundary layer resistance at feed side was increased along with the increasing of TEG

concentration due to the viscosity nature. The contribution of membrane resistance and

boundary layer resistance (both feed and permeate side) is expressed in Figure 4.11. During

increasing of TEG concentration (from 0 to 80%), the boundary layer resistance increased

from 21.9 to 132.2 x 105 m/s. Its contribution to the total resistance increased from 7.6 % to

33.2% which resulted in reducing of permeate flux. However, the boundary layer resistance

of SGMD is not a serious problem (Khayet et al., 2002; Khayet and Matsuura, 2011). It can

be recovered by rising the system with DI water and chemical agents.

Table 4.8 Membrane Resistance and Boundary Layer Resistance

TEG

Concentration

(%)

Permeate flux

(kg/m2.h)

MD

resistance

(105 m/s)

Boundary Layer resistance

(105 m/s)

0 3.14

266

21.9

10 2.61 26.3

20 2.34 29.4

30 2.09 32.9

40 1.83 37.6

60 1.57 43.8

70 1.08 63.7

80 0.52 132.3

69

Figure 4.11 Increasing of boundary layer resistance

The calculations of boundary layer resistance are expressed Appendix D

4.3.3 Bench scale with real TEG wastewater

Real wastewater was used to conduct the experiment on hollow fiber membrane after PWF

and synthetic TEG test. This experiment was aimed to evaluate the performance of SGMD

when operating with the real wastewater. Fouling analysis were conducted by checking the

PWF after cleaning the module with DI water and chemical agents. By operating the bench

scale hollow fiber module with real wastewater, it would be suitable to scale up the operating

condition for the pilot scale module, then designing of full scale plant. The operational

conditions of this experiment were feed flow rate of 2.4 L/min, feed temperature of 70oC

and sweeping gas velocity of 4.7 m/s.

4.3.3.1 Batch experiments

The real wastewater after treating by pre-treatment system was divided into 5 batches to

concentrate by SGMD system. The feed solution (real wastewater) was changed after each

batch without any system’s cleaning. By operating all experiments at the same condition

(sweeping gas flow rate of 0.255 L/min/fiber (4.7 m/s), feed temperature of 70oC and feed

flow rate of 2.4 L/min), the experimental results of Batch 1 and batch 5 are shown in Figure

4.12. After finishing 5 batches (total of 40 hours), fouling analysis was conducted to identify

different types of fouling.

50

60

70

80

90

100

0 10 20 30 40 60 70 80

Pro

po

rtio

n c

on

trib

uti

ng t

o r

esis

tan

ce

(%)

TEG concentration (%)

Membrane Resistance Boundary Layer Resistance

70

Figure 4.12 Experimental results of batch experiment with real wastewater.

It can be clearly observed (Figure 4.11) that the permeate flux of real wastewater of 5

batches were quiet similar with a small variation from 2.1 to 2.6 kg/m2.h. Overall, it was

stable at 2.4 kg/m2.h while that values for synthetic TEG (10%) and pure water were 2.61

and 3.14 kg/m2.h respectively. The lower permeate flux of real wastewater in comparison

with pure water and synthetic TEG can be explained by the presence of other impurities in

real wastewater. This such impurities would result in increasing feed boundary layer

resistance. All experimental data is shown in Appendix E

4.3.3.2 Continuously fed TEG wastewater (40 hours)

Figure 4.13 presents the experimental result of continuously-fed real wastewater. Due to

the effect of impurities, the boundary layer resistances were higher than that value of

synthetic TEG experiment. At first 16 hours, the permeate flux was about 2.4 kg/m2.h with

the TEG concentration less than 30 %. The flux reduced to 2.1 kg/m2.h at higher TEG

concentration after 16 hrs. It reached 1.6 kg/m2.h at the hour of 40th (final TEG concentration

was about 55 %). After finishing real wastewater experiment, the system was rinsed with

DI water and chemical agents to measure fouling resistance.

6

10

14

18

1.5

2.0

2.5

3.0

3.5

0 2 4 6 8 10

TE

G co

nce

ntr

atio

n (

%)

Per

mea

te f

lux

(k

g/m

2.h

)

Time (h)

Batch 1 Batch 5 Synthetic 10%

Pure water flux TEG concentration

71

Figure 4.13 Experimental result of continuously-fed real wastewater investigation

4.3.3.3 Continuously fed TEG wastewater (210 hours)

The real wastewater was continuously fill up to 10 L in the feed tank without draining any

concentrate. Figure 4.14 show the permeate flux of SGMD during 210 hours operating

continuously. This corresponds to 38 days of operation at 8 hours operation/day. It is clear

that the permeate flux was almost stable at 2.6 kg/m2.h within the first 60 hours. At the hour

in the range from 60 to 170, it reduced and stable at 2.3 kg/m2.h. After 170 hours of feeding

continuously, the permeate flux of membrane was stable at 1.5 kg/m2.h. Overall, the

permeate flux of SGMD after 210 hours operating continuously was reduced to 60% as

compared to the first hour. It proves that fouling phenomena is not a serious problem in MD

technology.

Figure 4.14 Permeate flux and TEG concentration during 210 hrs feeding

continuously

0

10

20

30

40

50

60

1.0

1.5

2.0

2.5

3.0

0 10 20 30 40 50

TE

G c

on

cen

trat

ion

(%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)

Permeate flux TEG concentration

0

20

40

60

80

100

0.0

0.5

1.0

1.5

2.0

2.5

3.0

0 40 80 120 160 200 240

TE

G c

once

ntr

atio

n (

%)

Per

mea

te f

lux (

kg/m

2.h

)

Time (h)

Permeate flux

TEG Concentration

72

From Figure 4.14, the permeate flux was about 0.1 kg/m2.h when TEG concentration

reached 98.7 %. It indicates that MD had the ability to concentrate non-volatile compound

until nearly 99%. All experimental data are presented in Appendix E

4.3.4 Fouling analysis

It can be clearly observed that the permeate flux achieved when investigating bench scale

module with real wastewater is lower than that value of synthetic TEG. There was some

impurities remained in pre-treated wastewater (95% turbidity and 66% oil and grease were

removed by pre-treatment system). The impurities in real wastewater (almost non-volatile

compounds such as metal ions, SS) accumulated on the feed surface of fiber which resulted

in pore clogging and increasing the thickness of boundary layer. DI water and chemical

agents (Citric and Oxalic acids) were used to clean the membrane module.

4.3.4.1 Batch experiment

After 5 batches (8 h/batch) of real wastewater experiment, fouling analysis was conducted.

In SGMD configuration, the boundary layer resistance at feed side is dominant for total

boundary layer (Khayet and Matsuura, 2011). In the total boundary layer resistance (29.26

x 105 m/s) of batch experiment, the contribution of feed boundary layer resistance and

permeate boundary layer resistance were 88.9% (26.03 x 105 m/s) and 11.1% (3.23 x 105

m/s) respectively. Figure 4.15 shows the contribution of different types of resistance to the

total resistance. After cleaning by chemical, the irreversible fouling resistance accounted for

2% of total resistance.

Table 4.9 Total Membrane Boundary Layer and Fouling Resistance Calculated from

Fouled Permeate Flux (batch operation)

Type of Resistance Value (105 m/s) Percentage

Total 384.9 100

Membrane 266 69.2

Boundary layers

(feed + permeate) 29.26 7.6

Fouling

- Recoverable

- Reversible

- Irreversible

65.45

16.17

7.85

17.0

4.2

2.0

73

Figure 4.15 Classification of types of resistances in SGMD batch operation

4.3.4.2 Continuously-fed experiment

Similar to fouling analysis of batch experiment, the membrane module was also cleaned by

DI water and chemicals after operating 40 hrs continuously. In Figure 4.16, the contribution

of boundary layer resistance (30.9%) into the total resistance was higher than that value in

batch experiment (23.2%)


Fouled Permeate Flux (continuously-fed operation)


Total 432.8 100

Membrane 266 61.5

Boundary layers


Fouling

- Recoverable

- Reversible

- Irreversible

109.7

16.17

7.85

25.3

3.7

1.8

Membrane

69.2 %

Boundary layer

7.6 %

Recoverable

17.0%

Reversible

4.2 %

Irreversible

2.0 %

Fouling

23.2 %

74

Figure 4.16 Resistance classification in SGMD continuously-fed real wastewater

Table 4.10 shows the comparison of resistance between batch and continuously-fed

experiment. Due to the increasing of TEG concentration and impurities, total resistance in

continuously-fed experiment (432.8 x 105 m/s) was higher than that value in batch

experiment (384.9 x 105 m/s)

Table 4.11 Comparison of Fouling and Other Resistance between Continuously-fed

and Batch Operation

Type of Resistance Value (105 m/s)

Continuously-fed Batch

Total 432.8 384.9

Membrane 266 266

Boundary layers


Fouling

- Recoverable

- Reversible

- Irreversible

109.7

16.17

7.85

65.45

16.17

7.85

From Figure 4.15 and 4.16, fouling resistance could be greatly removed by only using DI

water (bases on the percentage of recoverable fouling resistance in total fouling resistance).

Thus, it can be concluded that fouling in membrane distillation process is not a serious

operational problem. The bench scale hollow fiber membrane distillation, when operating

as SGMD configuration, has lower total resistance in comparison with DCMD

configuration. This is due to the fact that boundary layer resistance formed in SGMD was

very low (less than 10 times with that value in DCMD). The irreversible fouling resistance

in SGMD was higher than DCMD due to accumulating of air impurities (i.e. dust, oil

Membrane

61.5 %

Boundary layer

7.6 %

Recoverable

25.3%

Reversible

3.7 %

Irreversible

1.8 %

Fouling

30.9 %

75

compound from compressor) at the permeate side of membrane. It could not be removed

during cleaning procedure. For DCMD, the permeate side could be rinsed with DI water.

Table 4.12 Comparison of Fouling and Other Resistance between SGMD and DCMD

Bench Scale Hollow Fiber Membrane Distillation (Continuously-fed)

Type of Resistance

Value (105 m/s)

SGMD

(This study)

DCMD

(Yuthawong, 2014)

Total 432.8 730

Membrane 266 236

Boundary layers


Fouling

- Recoverable

- Reversible

- Irreversible

109.7

16.17

7.85

73

0

2.6

4.3.5 Evaluation of energy consumption of bench scale SGMD

To enhance the heat transfer efficiency, it was necessary to minimize the boundary layer

resistances (temperature polarization would be reduced consequently) (Criscuoli et al.,

2008). To maintain the driving force of SGMD process, the heat should be supplied into

feed solution and the sweeping gas should be pumped continuously into the hollow fiber.

These steps required energy consumption. In this study, the energy efficiency of operating

process was evaluated. This was the energy required to maintain the driving force in the

MD process after the system started.

There were 4 main components which consumed energy in bench scale SGMD system:

feed pump, heater, control box, and air compressor. The electricity consumptions of three

first components (feed pump, heater, control box) were measured by a power meter.

Meanwhile, the energy consumption of air compressor was calculated theoretically based

on amount of air delivered and the full load power of compressor. The calculations for

energy ratio are shown in Appendix F

76

4.3.5.1 Synthetic TEG experiments

Batch Experiments

Figure 4.17, 4.18 and 2.19 presents the ratio of energy consumption/Qp (kW/kg)

corresponding to TEG concentration during concentrating process.

Figure 4.17 Energy consumption/Qp during concentrating TEG at 10% initial

concentration

When concentration TEG from 10 to 30%, the energy consumption/Qp was about 1.2 kW/kg

at the first 8 hrs. During concentrating TEG process, the volume of liquid in heat tank

reduced. Thus, the heater consumed less energy. However, at higher TEG concentration

(>30%), the energy consumption of the heater was lower due to the high latent heat of TEG

Figure 4.18 Energy consumption/Qp during concentrating TEG at 30%

initial concentration

5

15

25

35

45

55

0.8

1.2

1.6

2.0

2.4

0 2 4 6 8 10 12 14

TE

G C

on

cen

trat

ion

(%

)

En

erg

y C

on

sum

pti

on

/Qp

(k

W/k

g)

Time (h)

Energy ratio TEG Concentration

25

35

45

55

65

0.8

1.2

1.6

2.0

2.4

0 2 4 6 8 10 12 14

TE

G C

once

ntr

atio

n (

%)

Ener

gy C

onsu

mpti

on/Q

p (

kW

/kg)

Time (h)


77

Figure 4.19 Energy consumption/Qp During concentrating TEG at 60% initial

concentration

The permeate flux and energy consumption had small variation in every each hour. Thus,

the energy ratio was not stable but it showed a trend. At low TEG concentration, permeate

flux was very high. Thus, the ratio of energy consumption (kW/kgh-1) was very low.

Increasing TEG concentration, the flux reduced and pump consumed more power. Along

with operating timeline, the volume of feed aqueous solution in feed tank reduced which

leads to reducing power consumption of the heater. To be concluded, when TEG

concentration was in the range from 10 to 60%, the ratio of energy consumption/Qp was

approximately 1.46 kW/kgh-1. At the later period, the average value of energy ratio was

~3.39 kW/kgh-1.

Continuously-fed experiments

Figure 4.20 Energy consumption/Qp during concentrating TEG for 40 hours

50

60

70

80

90

100

2

3

4

5

0 5 10 15 20 25

TE

G C

once

ntr

atio

n (

%)

Ener

gy C

onsu

mpti

on/Q

p

(kW

/kg)

Time (h)


5

15

25

35

45

55

1.0

1.2

1.4

1.6

1.8

0 5 10 15 20 25 30 35 40

TE

G c

on

centr

atio

n(%

)

En

ergy C

on

sum

pti

on/Q

p(k

W/k

g)

Time (h)

Energy ratio TEG concentration

78

The energy consumption, the average ratio of energy consumption/ permeate flow when

concentrating TEG from 10 to 50 % was about 1.29 kWh/kg. It means that SGMD

configuration consumed less energy than DCMD (3.55 kWh/kg) and higher than VMD (1.1

kWh/kg) (Criscuoli et al., 2008).

4.3.5.2 Real wastewater experiment

Batch experiments

Figure 4.21 shows the comparison of energy consumption/ permeate flow between pure

water, synthetic TEG and real waste water. Due to the highest permeate flux in pure water

test (comparing with other feed solutions), the ratio of energy consumption/Qp of PWF

was lowest (1.09 kW/kg) while that value for real wastewater and synthetic TEG were

1.45 and 1.24 kW/kg respectively.

Figure 4.21 Energy consumption/Qp of batch experiment with real wastewater

The permeate flux at hour of 7th and 8th h in batch 5 reduced significantly. Thus, the energy

ratio of those hours increased very high.


6

10

14

18

1.0

1.5

2.0

2.5

3.0

0 1 2 3 4 5 6 7 8T

EG

co

nce

ntr

atio

n (

%)

Ener

gy c

onsu

mpti

on/Q

p (

kW

/kg)

Time (h)

Batch 1 Batch 5

Synthetic 10% Pure water flux

TEG concentration

79

Figure 4.22 Energy consumption/Qp of continuously-fed experiments with real

wastewater

The average value of energy consumption/Qp during 40 hours of operating was 1.35 kW/kg

(TEG concentration increased from 10 to 43%). This ratio is lower than that in batch

experiment (1.45 kW/kg). As presented in Figure 4. 14, the average energy consumption/Qp

of the period of first 160 hrs (when TEG concentration achieved 98.7%) was about 2.67

kWh/kg. When TEG concentration increased to higher than 50%, the energy consumption

was similar with the earlier period (low than 50%) but the permeate flux reduced. Thus, the

ratio of energy consumption/Qp increased.

4.4 Optimum Condition of Pilot Scale SGMD and Evaluate Energy Consumption

The experimental procedure in pilot scale SGMD unit was similar with that in bench scale

SGMD study. The system was first conducted with verification process. In this step, pure

water flux and rejection test were investigated respectively. Later on, synthetic TEG and

real wastewater were used to determine the system’s performance. For each experiment,

energy consumption evaluation and membrane fouling analysis (with the same cleaning

procedure as bench scale unit) were implemented. Last, the financial analysis was

evaluated for this pilot scale unit.

4.4.1 System investigation

4.4.1.1 Pure water flux

In this section, effect of gas velocity was investigated clearly by varying the gas inlet

velocity at many values (4.29, 4.98, 5.25, 6.06, 6.76, 7.57, 8.07, 8.58, 9.09, 9.59, 10.09,

10.60 m/s). According to section 4.3, the optimum feed temperature was 70oC. This feed

temperature was applied to pilot scale study. Besides, It was also proved that the feed flow

rate had negative effect on SGMD performance. Thus, the feed pump of pilot scale system

0

10

20

30

40

50

0.0

0.5

1.0

1.5

2.0

0 5 10 15 20 25 30 35 40

TE

G c

on

cen

trat

ion

(%)

En

ergy c

on

sum

pti

on

/Qp

(kW

/kg)

Time (h)


80

was operated at normal capacity (60L/min). Table 4.13 shows the experimental results of

pure water flux (PWF) investigation on pilot scale module.

The highest permeate flux that could be achieved in pilot scale was 1.99 kg/m2.h at the gas

velocity inside each fiber of 10.09 m/s (gas inlet flow rate of 0.57 L/min.fiber). For the

bench scale module, that value was 3.14 kg/m2.h at the gas velocity inside each fiber of 4.7

m/s (gas inlet flow rate of 0.255 L/min.fiber). It could be concluded that the fiber’s length

had a significant effect on SGMD process. The temperature profile of sweeping gas relied

along the length. The relationship between PWF and the ratio of energy consumption/Qp is

simulated in Figure 4.23. At higher gas velocity, the fiber has a greater opportunity of be

wetted. The hydrostatic pressure which was created by high sweeping gas velocity could

suck the feed liquid througout the membrane pores. The local permeate flux 𝐽𝑖(𝑥) along the

module length (L) can be calculated by using Equation 4.5 (Khayet and Matsuura, 2011).

𝐽𝑖 =1

𝐿∫ 𝐽𝑖(𝑥). 𝑑𝑥

𝐿

𝑜 (4.5)

The temperature varied along the length of fiber. The local temperature depends strongly on

the gas inlet flow rate. Thus, it is necessary to determine the profile of temperature and

vapour pressure along the membrane module length. Then, the local driving force would be

evaluated. The optimum condition was decided at the gas velocity of 8.07 m/s (gas inlet

flow rate of 0.44 L/min.fiber), feed inlet temperature of 70oC. Feed flow rate and sweeping

gas temperature have less effect. The flux achieved at this conditions was 1.94 kg/m2.h. At

this operating condition, the pressure at both feed side and permeate side was kept constant

at 0.4 bar. All experiment data is provided in Appendix E

Figure 4.23 Pure water flux at different sweeping gas inlet flow rate of pilot scale

study

0.3

0.4

0.5

0.6

0.7

0.8

0.5

1.0

1.5

2.0

2.5

4.0 5.0 6.0 7.0 8.0 9.0 10.0 11.0

En

ergy c

onsu

mpti

on (

kW

/kg)

Per

mea

te f

lux (

kg/m

2.h

)

Gas inlet flow rate per fiber (L/min)

81

Table 4.13 Experimental Results of PWF Investigation on Pilot Scale Module

Gas inlet

pressure

(bar)

Gas flow rate

(L/min)

Gas velocity

(m/s)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

0.10 70.90 4.29 0.98 0.58

0.15 82.20 4.98 1.19 0.55

0.20 86.60 5.25 1.53 0.46

0.25 100.0 6.06 1.63 0.48

0.30 111.6 6.76 1.79 0.48

0.35 125.0 7.57 1.82 0.52

0.40 133.3 8.07 1.94 0.51

0.45 141.6 8.58 1.93 0.55

0.50 150.0 9.09 1.89 0.57

0.55 158.3 9.59 1.95 0.59

0.60 166.6 10.09 1.99 0.60

0.65 175.0 10.60 1.99 0.63

Yuthawong (2014) pointed out that the sweeping gas inlet temperature has nearly negligible

effect on permeate flux. This is due to the low thermal conductivity of ambient air. At any

inlet temperature, the sweeping gas heated up very fast and achieved nearly the feed

temperature as soon as the entrance position of fiber. Besides, the contact time between

liquid flow and gas flow was longer than bench scale module due to pilot scale used higher

fiber’s length (nearly double length). It could be understood that almost local positions at

permeate side of membrane had the similar temperature. Other researches about the effect

of sweeping gas inlet temperature also had the same conclusion (Lee and Hong 2001; Xie

et al., 2009)

Table 4.14 Membrane Resistance and Membrane Coefficient

Membrane Mass transfer

mechanism

Yuthawong

(2014)

This study

Hollow fiber

SGMD

(0.45 µm)

Knudsen and

molecular diffusion

combined

Bw

(10-7 m/s)

Rw

(105 m/s)

Bw

(10-7 m/s)

Rw

(105 m/s)

0.25 406 0.23 448

According to Table 4.14, the membrane resistance (Rw) in this study was lower than that

value which was studied by Yuthawong (2014). It could be due to using an used membrane

module to start this study. This membrane surface had already been affected by the

irreversible fouling in the past. This slight higher resistance would lead to lower membrane

82

permeability. However, this issue was overcome by operating the module at optimum

condition.

4.4.1.2 Rejection

Before investigating with synthetic TEG and real wastewater, the rejection test were

conducted on pilot MD module. The goal of this experiment was to ensure that the system

was working under membrane distillation process. In bench scale study, salt solution 1%

was used. However, copper in heat exchanged could potentially be corroded

by salt. Thus, TEG 10% was used in this experiment. The intention of rejection test was still

remained the same: separate between volatile and non-volatile compound. During 6 hours

of experiment, TEG concentration in feed tank increased leading to reducing permeate flux

from 1.91 to 1.57 kg/m2.h. However, the rejection of MD was still 99.99% and remained

the same in the whole process.

Figure 4.24 Rejection results for pilot scale SGMD process of pilot scale study

It can be clearly observed from Figure 4.24 that the pilot scale module worked as membrane

distillation at the optimum condition and only volatile compound (water vapor) passed

through membrane pores.

4.4.2 Synthetic TEG

The performance of pilot scale system was investigated with synthetic TEG at two level of

sweeping gas velocity (2.1 and 8.07 m/s). The reason for operating at different gas flow rate

was to have a deeper understanding on the effect of gas velocity on MD process. In this

experiment, the effect of boundary layer resistance was evaluated (it theoretically creates

less effect in SGMD configuration (Khayet and Matsuura, 2011). All experiments were

conducted at feed temperature of 70oC, feed flow rate of 60 L/min with the duration of 8 hrs

per experiment. Initially, the feed tank was filled up to 52 L by synthetic TEG (preparing

by mixing DI water with stock TEG 90%). All experiment data is provided in Appendix E

0

20

40

60

80

100

1.4

1.6

1.8

2.0

0 1 2 3 4 5 6 7

Rej

ecti

on (

%)

Per

mea

te f

lux

(kg

/m2.h

)

Time (h)

Permeate flux Rejection

83

4.4.2.1 Batch Experiment

Low gas velocity experiment (2.1 m/s)

At the first stage, the pilot scale MD system was operated with the condition obtains: feed

flow rate of 60 L/min, feed temperature of 70oC and sweeping gas velocity of 2.1 m/s (flow

rate of 34.5 L/min). the results would contribute to the understanding of effect of sweeping

gas inlet flow rate on SGMD performance. The concentration of TEG in synthetic

wastewater varied from 10 to 80%.

Figure 4.25 Permeate flux and TEG concentration synthetic TEG 10% initial




8

10

12

14

16

0.6

0.7

0.8

0.9

1 2 3 4 5 6 7 8

TE

G C

on

centr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


20

30

40

50

60

0.6

0.7

0.8

0.9

1.0

1 2 3 4 5 6 7 8

TE

G C

once

ntr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


84

Figure 4.26 represents the relation between permeate flux and TEG concentration during

concentrating process. The flux at first three hours was not stable. This is normal because

normally, the system needed to spend few hours since the experiment started-up to reach

the stable operating condition. The average permeate flux of first 6 hours was 0.8 kg/m2.h

while TEG concentration kept increasing from 25 to 35%. Afterward, the flux reduced

slightly to 0.7 kg/m2.h.





As seen from Figure 4.25 to 4.29, when TEG concentration in feed solution increased from

10 to 60%, permeate flux was almost stable at 0.8 kg/m2.h. After that, the permeate flux

reduced continuously from 0.8 to 0.4 kg/m2.h when TEG concentration reached 90%.

Overall, the permeate flux only decreased a half (50%) during concentrating TEG in

synthetic wastewater from 10 to 90%. During the period of concentrating TEG from 10 to

40

50

60

70

80

0.5

0.6

0.7

0.8

0.9

1 2 3 4 5 6 7 8

TE

G C

once

ntr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


60

70

80

90

100

0.4

0.5

0.6

0.7

1 2 3 4 5 6 7 8

TE

G C

once

ntr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


85

80% (about 50 first hours), the ratio of energy consumption/Qp was about 0.6 kW/kg. At

higher TEG concentration, it rapidly increased (7 kW/kg at TEG concentration of 99%).



Optimum gas velocity experiment (8.07 m/s)

In this scenario, the pilot scale MD system was investigated with the optimum gas velocity

(8.07 m/s). This condition was selected from the pure water test (balancing between

permeate flux and energy consumption). Synthetic TEG was used as feed solution at

different initial concentration (by mixing stock TEG 90% with DI water). Due to effect of

fiber’s length, the results from this experiment showed the limitation of pilot scale hollow

fiber module in comparison with bench scale experiments. Figure 4.30 presents the

experimental results of first synthetic TEG experiment at 10% initial concentration.



75

80

85

90

95

100

0.0

0.2

0.4

0.6

0.8

1 3 5 7 9 11 13

TE

G C

once

ntr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


y = -0.0227x + 1.8363

R² = 0.4442

y = 1.4933x + 8.4478

R² = 0.96379

12

15

18

21

24

1.0

1.2

1.4

1.6

1.8

2.0

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


86

At the first hours of operating (initial TEG concentration of 10%), the permeate flux

achieved was 1.9 kg/m2.h. However, the permeate flux reduced in the following hours due

to the effect of concentration polarization. It decreased from 1.91 to 1.7 kg/m2.h, after 8 hrs

operating continuously. At the 8th hours, TEG concentration reached 21%. Implying 50%

of the volume was removed.



At the second batch, TEG concentration at the initial hour was similar with the final

concentration Figure 4.31 (about 20%). However, the effect of concentration polarization

was very less at first hour. Thus, the flux still reached 1.8 kg/m2.h. At the 8th hours, TEG

concentration reached 41 % with the flux of 1.51 kg/m2.h. Once the concentration increased,

the effect of concentration polarization became more evident. The latent heat of TEG

compound is higher than that of water. Thus, the surface temperature at the feed side

increased. Besides, thermal conductivity also increased along with increasing TEG

concentration (effect of viscosity).

y = -0.03x + 1.7824

R² = 0.5968y = 2.7106x + 17.356

R² = 0.971315

20

25

30

35

40

45

1.0

1.2

1.4

1.6

1.8

2.0

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion

(%

)

Per

mea

te f

lux (

kg/m

2.h

)

Time (h)


87





When TEG concentration increased from 30 to 70%, permeate flux reduced slightly from

1.6 to 1.2 kg/m2.h. as seen in Figure 4.32. But while increasing the initial TEG concentration

to 40% and 60%, the flux was strongly affected by TEG concentration due to its viscosity

nature as presented in Figure 4.33. After the 6th hour the flux decline could be observed.

y = -0.0327x + 1.7121

R² = 0.729y = 3.6805x + 26.666

R² = 0.977425

35

45

55

65

1.0

1.2

1.4

1.6

1.8

2.0

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion

(%

)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


y = -0.0381x + 1.59

R² = 0.9083

y = 4.0748x + 36.899

R² = 0.985335

45

55

65

75

1.0

1.2

1.4

1.6

1.8

1 2 3 4 5 6 7 8

TE

G c

once

ntr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


88



When the TEG concentration increased from 60 to 70 % (Figure 4.34), permeate flux was

about 1.15 kg/m2.h. At higher TEG concentration (>70%), permeate flux decreased

dramatically to less than 0.4 kg/m2.h at TEG concentration of 88%. At this situation, MD

was strongly affected by the feed viscosity.

4.4.2.2 Continuously-fed synthetic TEG

Synthetic TEG (10%) was used as the feed solution in this experiment. The feed tank was

filled up every 8 hours by synthetic TEG 10% solution and TEG samples were collected.

As presented in Figure 4.35 while feeding synthetic TEG 10%, the concentration decreased

significantly and the permeate flux increased consequently. However, its concentration

increased in whole process (from 10% to 65%) due to the continuous feeding progress.

Overall, the permeate flux was stable around 1.8 kg/m2.h when the TEG concentration was

less than 30 %. When the TEG concentration was in the range from 35% to 50%, permeate

flux was about 1.6 kg/m2.h. At higher TEG concentration (>50%), the flux reduced to less

than 1.2 kg/m2.h.

4.4.2.3 Boundary layer resistance

In synthetic TEG experiment, the permeability of MD process was not affected by any

impurities. The reduction of permeate flux was dependant on the increasing of boundary

layer resistances. Membrane resistance at the optimum condition was calculated as 448 x

105 m/s. The overall membrane resistance and boundary layer resistance from TEG

concentration of 0 - 80 % is presented in Table 4.15.

y = -0.0928x + 1.422

R² = 0.7082y = 3.6961x + 58.848

R² = 0.998150

60

70

80

90

100

0.2

0.6

1.0

1.4

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion

(%

)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


89

Figure 4.35 Experimental result of continuously-fed synthetic TEG investigation of

pilot scale study

Table 4.15 Membrane Resistance and Boundary Layer Resistance

TEG

Concentration

(%)

Permeate flux

(kg/m2.h)

MD resistance

(105 m/s)

Boundary Layer resistance

(105 m/s)

0 1.95

448

51.8

10 1.91 52.9

20 1.87 54.1

30 1.67 60.5

40 1.51 66.9

60 1.10 91.9

70 0.99 102.1

80 0.76 133.0

Similar with SGMD in bench scale study, the boundary layer resistance at feed side

increased along with the increase in TEG concentration due to its viscosity. The contribution

of membrane resistance and boundary layer resistance is expressed in Figure 4.29. As the

TEG concentration increased from 0 to 80%, the boundary layer resistance increased from

51.8 to 133 x 105 m/s. Its contribution to the total resistance increased from 10.4 % to 22.5%

which was the main reason for the reduction of permeate flux. However, the boundary layer

resistance of SGMD is not a fatal problem to the membrane itself. (Khayet et al., 2002;

Khayet and Matsuura, 2011). It can be removed by rising the system with DI water and

chemical agents.

5

15

25

35

45

55

65

75

0.8

1.0

1.2

1.4

1.6

1.8

2.0

0 5 10 15 20 25 30 35 40 45

TE

G c

on

cen

trat

ion

(%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


90

Figure 4.36 Proportion of boundary layer resistance and membrane resistance of

pilot scale study

4.4.3 Real TEG wastewater

The real wastewater experiment aimed to evaluate the performance of SGMD configuration

comparing with PWF and synthetic TEG test. Together, fouling analysis was conducted by

checking the PWF after cleaning the module with DI water and chemical agents. All

experiment results of pilot scale study with real wastewater were used to design the full

scale plant. The operational conditions of this experiment were feed flow rate of 60 L/min,

feed temperature of 70oC and sweeping gas velocity of 8.07 m/s. At the beginning, similar

to synthetic TEG experiment, the feed tank was filled up by wastewater at 52 L. Each

experiment was investigated with 8 hours of operating the system. This 8-hours represented

for the national working time per day of many countries. All experimental data is provided

in Appendix E

4.4.3.1 Batch operations

The pre-treated real wastewater was divided into 5 batches to conduct experiment on SGMD

system. The feed solution (real wastewater) was replaced after each batch without any

cleaning. All batches were operated at the same condition (sweeping gas flow rate of 133.3

L/min.fiber (8.07 m/s), feed temperature of 70oC and feed flow rate of 60 L/min).The

experimental results of Batch 1 and batch 5 are shown in Figure 4.37. Fouling analysis was

conducted to identify different types of fouling after finishing 5 batches (total of 40 hours).

50

60

70

80

90

100

0 10 20 30 40 60 70 80

Res

ista

nce

s in

MD

(%

)

TEG concentration (%)

Membrane Resistance Boundary Layer Resistance

91

Figure 4.37 Experimental results of batch experiment with real wastewater of pilot

scale study

From Figure 4.37, the permeate flux of real wastewater observed from batch 1st were very

different with batch 5th. At the first batch, permeate flux was almost stable about 1.8 kg/m2.h.

However, the flux achieved in batch 5th was lower (1.7 at the first hour and 0.8 at the 8th

hour). While that values for synthetic TEG (10%) and pure water were 1.95 and 1.91 kg/m2.h

respectively. The lower permeate flux of real wastewater in comparison with pure water and

synthetic TEG can be explained by the presence of other impurities in real wastewater. This

such impurities would result in increasing feed boundary layer resistance and boundary layer

resistance.

4.4.3.2 Fouling analysis for batch operation

After 5 batches (8 hrs/batch) of real wastewater experiment, fouling analysis was conducted

to evaluate the recoverable ability of membrane. In SGMD process, the feed boundary layer

resistance is dominant for total boundary layer (Khayet and Matsuura, 2011). In the total

boundary layer resistance (56.5 x 105 m/s), feed boundary layer resistance and permeate

boundary layer resistance contributed 46.1 x 105 m/s and 10.4 x 105 m/s respectively. Figure

4.43 shows the contribution of different types of resistance to the total resistance. After

cleaning by chemical, the irreversible fouling resistance accounted for 1.72% of total

resistance.

4

8

12

16

20

24

0.0

0.5

1.0

1.5

2.0

2.5

0 1 2 3 4 5 6 7 8

TE

G co

nce

ntr

atio

n (

%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)

Batch 1 Batch 5Synthetic 10% Pure water fluxTEG concentration

92


Fouled Permeate Flux


Total 544.7 100

Membrane 448.2 82.3

Boundary layers


Fouling

- Recoverable

- Reversible

- Irreversible

43.4

7.3

9.4

4.3

1.3

1.7

Figure 4.38 Classification of types of resistances in SGMD batch operation of pilot

scale study

4.3.3.3 Fed-continuously experiment 40 hours

Fed-continuously experiment 40 hours

Figure 4.39 presents the experimental result of continuously-fed real wastewater on pilot

scale SGMD module. Due to the effect of impurities (i.e. SS, metal), the boundary layer

resistances were higher than that value of synthetic TEG experiment. The first hour, MD

achieved permeate flux of 1.9 kg/m2.h. At first 5 hours, the permeate flux was higher than

1.7 kg/m2.h with the TEG concentration less than 15 %. The flux reduced to the range

between 1.4 and 1.65 kg/m2.h at higher TEG concentration from the hour of 5th and 28th

(TEG concentration less than50%). It reached 0.9 kg/m2.h at the hour of 40th (final TEG

concentration was about 69.4 %). After finishing real wastewater experiment, the system

Membrane

82.3 %

Boundary layer

10.4 %

Recoverable

4.3 %

Reversible

1.3 %

Irreversible

1.7 %Fouling

7.4 %

93

was rinsed with DI water and chemical agents to measure all types of resistance (fouling

analysis).

Figure 4.39 Experimental result of continuously-fed real wastewater investigation of

pilot scale study (40 hours)

Fed-continuously experiment 72 hours

The concentrated wastewater from previous experiments (TEG concentration was in the

range from 20-80%) was restored and mixed with the remain wastewater in the storage

tank. Therefore, after mixing, the TEG concentration of mixed liquid was 19.6 %.

The continuously-fed experiment 72 hours was conducted to evaluate both MD

performance and to know the concentrating ability of MD process. Figure 4.40 shows the

relation between permeate flux and TEG concentration during 72 hours of operating. For

the experiment using real wastewater on the clean MD module, both batch test and

continuously test, the permeate flux of first hour was very high (about 1.8 kg/m2.h).

However, the initial TEG concentration in this experiment was higher than that in the 40-

hours continuously-fed experiment. As it can be observed from Figure 4.40, the average

permeate flux when TEG concentration less than 60% was 1.6 kg/m2.h. Later on, the

average permeate flux reduce to 1 kg/m2.h once the TEG concentration was in the range

from 60-90%. The flux achieved was 0.4 kg/m2 when TEG concentration reached 99.8%.

5

15

25

35

45

55

65

75

0.8

1.0

1.2

1.4

1.6

1.8

2.0

0 5 10 15 20 25 30 35 40 45

TE

G c

on

cen

trat

ion

(%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)Permeate flux TEG concentration

94

Figure 4.40 Experimental result of continuously-fed real wastewater investigation of

pilot scale study (72 hours)

4.3.3.3 Fouling analysis for continuously-fed operation

The impurities in real wastewater (almost non-volatile compounds such as metal ions, SS)

accumulated on the feed surface of fiber which resulted in pore clogging and increasing the

boundary layer resistance. DI water and chemical agents (Citric and Oxalic acids) were used

to clean the membrane module. Figure 4.41 presents the contribution of different types of

resistance. After cleaning with chemical agents, irreversible fouling resistance remained

about 9.4 x 105 m/s (1.3% of total resistance). Table 4.17 shows the typical value of all

fouling resistances.

Table 4.17 Total membrane boundary layer and fouling resistance calculated from

fouled permeate flux (continuously-fed operation)


Total 744.4 100

Membrane 448.2 60.2

Boundary layers


Fouling

- Recoverable

- Reversible

- Irreversible

181.6

27.9

9.4

24.4

3.8

1.3

0

10

20

30

40

50

60

70

80

90

100

0.0

0.4

0.8

1.2

1.6

2.0

0 10 20 30 40 50 60 70 80

TE

G c

on

cen

trat

ion

(%)

Per

mea

te f

lux

(kg/m

2.h

)

Time (h)


95

Figure 4.41 Classification of Resistances in SGMD Continuously-fed Real

Wastewater of Pilot Scale Study

Table 4.18 shows the comparison of resistance between batch and continuously-fed

experiment of pilot scale MD module. Due to the increasing of TEG concentration and

impurities, total resistance in continuously-fed experiment (744.4 x 105 m/s) was higher than

that value in batch experiment (544.7 x 105 m/s). During concentrating continuously,

impurities in real wastewater attached on the membrane surface without removing. Thus,

fouling resistance in continuously-fed experiment was more serious than batch experiment.

Table 4.18 Comparison of Fouling and Other Resistance between Continuously-fed

and Batch Operation


Continuously-fed Batch

Total 744.4 544.7

Membrane 448.2 448.2

Boundary layers


Fouling

- Recoverable

- Reversible

- Irreversible

181.6

27.9

9.4

43.4

7.3

9.4

A comparison of fouling analysis results of pilot scale MD module is presented in Table

4.19. In the research of Yuthawong (2014), MD module was investigated at the non-optimal

condition. A high total resistance (1960 x 105 m/s) resulted in low permeate flux. Boundary

layer resistance of membrane module in this situation was high (963 x 105 m/s). Moreover,

Membrane

60.2 %

Boundary layer

10.4 %

Recoverable

24.4 %

Reversible

3.8 %

Irreversible

1.3 %

Fouling

29.4 %

96

due to the unsuitable cleaning procedure (using hydrochloric acid and sodium hydroxide),

the irreversible fouling resistance (67 x 105 m/s) remained after cleaning was higher than

this study (9.4 x 105 m/s). By using proper cleaning agents (citric and oxalic acid), the

irreversible fouling resistance was very less.

Table 4.19 Comparison of Fouling and Other Resistance between Normal Condition

and Optimum Condition of SGMD Hollow Fiber Membrane Distillation

(Continuously-fed)


This study Yuthawong (2014)

Total 744.4 1960

Membrane 448.2 406

Boundary layers

(feed + permeate) 77.2 963

Fouling

- Recoverable

- Reversible

- Irreversible

181.6

27.9

9.4

501

27

67

In this study, all calculations related to feed velocity (represented by feed flow rate) used

the cross flow velocity. This velocity was determined based on the feed flow rate and cross

sectional area of the liquid flow channel. However, the local velocity at different positions

was totally different due to the sharp of module. Figure 4.42 and 4.43 show the typical

velocities profile along the module. The flow simulation were done in Solidworks software.

The characteristic of flow affected mainly on membrane surface fouling. At the area of

turbulent flow, the fouling resistance was less than that at laminar flow.

97

Figure 4.42 Inlet velocity distribution in hollow fiber module

Furthermore, at the entrance area of liquid flow, the hollow fiber membrane had higher risk

of breaking due to exposing directly with the high liquid flow velocity. Those hollow fiber

membrane also acted as a barrier of entrance flow which led to reduce liquid velocity in

whole module and created the reverse velocity at the opposite side of entrance region.

Further information of flow simulation is shown in Appendix D

4.4.5 Evaluation of energy consumption of pilot scale SGMD

In general energy ratio is affected by two components: Energy consumption of whole system

(kW/h) and permeate flow rate (kg/h). It was different from bench scale SGMD system, the

pilot scale SGMD system used a heat exchanger to maintain the desired feed temperature.

98

Figure 4.43 Local Velocities inside membrane module

4.4.5.1 Synthetic TEG experiments

Batch experiments

Fingure 4.44, 4.45, 4.46, 4.47 and 4.48 show the ratio of energy consumption/Qp (kW/kg)

corresponding to TEG concentration during concentrating process. These were batch

experiments at different initial TEG concentration (10, 20, 30, 40, and 60%)

99



During operational process, the ratio of energy consumption/Qp was slightly increased from

0.53 to 0.58 kW/kg when TEG concentration went up from 10.8 to 19%. This was due to

the reduction of permeate flux during the process. At the 8th hour, the energy consumption

(kW/h) reduced significantly due to the small concentration feed volume. The heater would

consume less energy. This led to lower energy ratio as observed.



The necessary energy ratio to concentrate TEG from 20 to 35% was 0.58 kW/kg. This ratio

was 0.64 kW/kg when TEG concentration reached 40.8 %. This was due to the reduction of

permeate flux during increasing the viscosity in feed solution.

y = 0.0065x + 0.5359

R² = 0.4358

y = 1.4794x + 8.4896

R² = 0.9679

12

15

18

21

24

0.40

0.45

0.50

0.55

0.60

0.65

0.70

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion

(%

)

Ener

gy c

onsu

mpti

on/Q

p (

kW

/kg)

Time (h)


y = 0.0107x + 0.5448

R² = 0.6661y = 2.7106x + 17.356

R² = 0.971315

20

25

30

35

40

45

0.40

0.45

0.50

0.55

0.60

0.65

0.70

1 2 3 4 5 6 7 8

TE

G c

once

ntr

atio

n (

%)

En

ergy c

onsu

mpti

on/Q

p (

kW

/kg)

Time (h)


100



Similar to the previous batch (20% initial TEG concentration), the necessary energy ratio

was 0.58 kW/kg when TEG concentration increased from 30 to 35%. The trend of energy

ratio kept on increasing after that. To concentrate TEG from 35 to 50%, the average energy

ratio required was 0.63%.



y = 0.0135x + 0.5598

R² = 0.8177

y = 3.6805x + 26.666

R² = 0.977430

35

40

45

50

55

60

0.40

0.45

0.50

0.55

0.60

0.65

0.70

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion

(%

)

En

ergy c

on

sum

pti

on

/Qp

(kW

/kg)

Time (h)


y = 0.016x + 0.6052

R² = 0.8941y = 4.0748x + 36.899

R² = 0.985335

45

55

65

75

0.50

0.55

0.60

0.65

0.70

0.75

0.80

1 2 3 4 5 6 7 8

TE

G c

once

ntr

atio

n (

%)

En

ergy c

onsu

mpti

on/Q

p (

kW

/kg)

Time (h)


101



It could be clearly observed from Figure 4.47 and 4.48, the ratio of energy consumption/Qp

increased together with the increasing of TEG concentration. This ratio was depended

strongly on both energy consumption (kW/h) of the system and permeate flow (kg/h). To

concentrate TEG from 50 to 60%, the average energy ratio required was 0.79 kW/kg. When

TEG concentration reached 88%, SGMD process required 2.1 kW/kg to maintain the driving

force.


Different from batch experiment, the permeate flux in continuously-fed experiment was

lower due to the continuing built up of fouling resistance. However, the feed solution already

owned some heat when the feed tank was filled up. As a result, the energy consumption per

hour was reduced. Therefore, the energy ratio was not so different from that values in batch

experiment. When concentrating TEG from 10 to 35 %, the value of energy ratio was 0.56

kW/kg. It required 0.94 kW/kg when TEG concentration reached 65 %.

y = 0.1395x + 0.4186

R² = 0.5504

y = 3.6961x + 58.848

R² = 0.9981

50

60

70

80

90

100

0.7

1.1

1.5

1.9

2.3

1 2 3 4 5 6 7 8

TE

G c

on

cen

trat

ion

(%

)

En

ergy c

on

sum

pti

on

/Qp

(kW

/kg)

Time (h)


102

Figure 4.49 Energy consumption/Qp during concentrating TEG for 40 hours of pilot

scale study

4.4.5.2 Real wastewater experiment

Batch experiments

Figure 4.50 presents the ratio of energy consumption/ permeate flow between pure water,

synthetic TEG and real waste water. In pure water test, the flux achieved was higher than

in synthetic and real wastewater. Therefore, the average energy ratio of pure water test

(within 8 hours) was 0.51 kW/kg while that value were 0.54 and 0.57 kW/kg for real

wastewater and synthetic TEG respectively.

Figure 4.50 Energy consumption/Qp of batch experiment with real wastewater of

pilot scale study

5

15

25

35

45

55

65

75

0.4

0.6

0.8

1.0

0 5 10 15 20 25 30 35 40 45

TE

G c

on

cen

trat

ion

(%)

Ener

gy c

onsu

mpti

on/Q

p (

kW

/kg)

Time (h)Energy ratio TEG concentration

4

8

12

16

20

24

0.40

0.45

0.50

0.55

0.60

0.65

0.70

1 2 3 4 5 6 7 8

TE

G co

nce

ntr

atio

n (

%)

En

ergy c

on

sum

pti

on

/Qp (

kW

/kg)

Time (h)

Real wastewater Synthetic 10%

Pure water TEG concentration

103


Figure 4.51 shows the energy ratio of continuously-fed experiment (real wastewater test).

c

Figure 4.51 Energy consumption/Qp of continuously-fed experiments with real

wastewater of pilot scale study

In bench scale, the feed temperature was maintained by the electric heater (contact directly

to the feed tank). Thus, the heat generated by electricity was used to maintain the feed

temperature. Meanwhile, the feed temperature in pilot scale system was maintained by a

heat exchanger.

In pure water test, the ratio of energy consumption/Qp was 1.09 kW/kg in bench scale study

(PWF of 3.14 kg/m2.h). This value was 0.51 kW/kg in pilot scale study (PWF of 1.94

kg/m2.h). Those values indicated that it is more economic when using heat exchanger to

maintain the driving force of SGMD process. Thus, optimization of MD (system design

and economic analysis) is very necessary and urgent to compare with other existing

technology. MD process should become possible for industrial application. Some strategies

have been applied on MD process for energy saving. Those are: well-designed system with

low effect from temperature polarization and concentration polarization, minimized heat

loss (reduce membrane thickness), heat recovery (by using heat exchanger). The ultimate

goal of optimizing process is to enhance MD process productivity.

4.4.5.3 Energy analysis

In technology innovation, energy consumption of MD process is an important factor for this

energy intensive technology. The components consuming energy in MD process are:

thermal energy to heat the feed liquid, energy required for circulation pump and compressor.

The heat energy requirement is more than 90% of the total energy consumption of MD

5

15

25

35

45

55

65

75

0.4

0.6

0.8

1.0

0 5 10 15 20 25 30 35 40 45

TE

G c

on

cen

trat

ion

(%)

En

ergy c

on

sum

pti

on

/Qp

(kW

/kg)

Time (h)Energy ratio TEG concentration

104

process (Khayet and Matsuura, 2011). The ratio of energy consumption/Qp of MD process

at different TEG concentration of the continuously-fed experiment is presented in Figure

4.49. During concentrating TEG from 9.69 to 50 %, the value of energy consumption/Qp

was about 0.7 kW/kg. The permeate flux reduced when increasing TEG concentration. Thus,

the value of energy consumption/Qp increased during concentrating process

The thermal efficiency in SGMD is very high (93.2%) while that value for DCMD is

between 50-80% (Khayet and Matsuura, 2011). Thus, it can be concluded that the heat loss

is not significant in MD process. The average energy required to concentrate TEG in real

wastewater from 10 to 60 % is 0.624 kW/kg (624 kW/m3). In literature energy consumption

was evaluated for DCMD and VMD at lab scale (Criscuoli et al., 2008). The authors

conducted a study on plate-and- frame membrane module (membrane area of 40 cm2). The

lowest ratio of energy consumption/Qp obtained were 3546.3 kW/m3 and 1108.4 kW/m3 in

DCMD and VMD respectively. The performance of VMD was proven to be better than

DCMD (based on permeate flux, energy consumption and thermal efficiency). As yet,

almost studies were done theoretical modeling. Thus, it is very important to conduct study

on realistic energy consumption. Additional information about heat balances are provided

in Appendix D

4.5 Full Scale SGMD Plant Design

Based on the results of this study, a full scale SGMD system was designed for application

at the Gas Separation Plant. The main desired design criteria for full scale SGMD plant was

to concentrate 1 m3 to 0.2 m3 per day (TEG concentration increased from 10% to 50%). All

the costs (investment, operation and maintenance) was estimated based on the market price

at the designing stage. Operational condition was the optimum conditions from this study.

The details of calculations and CAD drawings are provided in Appendix B

4.5.1 Overall system design

Real TEG wastewater from PTT gas separation plant was characterized in order to evaluate

the degree of pre-treatment required prior to membrane distillation (Imdakm and Matsuura)

unit. Thus, proper pretreated wastewater was further efficiently treated by the MD process.

The result of wastewater, volatile compound and heavy metal analyzed by ALS Company

are shown in Table 4.1. A pre-treatment system is a need to keep the feed to the MD

relatively clean so that long term effective operation could be made possible. It can be noted

that suspended solids (SS) and oil and grease (O&G) have to be eliminated as they can

potentially damage membrane distillation unit. Another issue would be iron concentration

observed in sample 2 as 41 mg/L. As iron can precipitate in membrane distillation system

when the concentration of iron reaches its solubility limit as TEG concentrates in the MD

unit. However, it seemed that high iron concentration was not consistent as in sample 3, iron

concentration was very low (1.2 mg/L) and it might be due to process instability at the GSP.

The third issue that the pre-treatment system needs to address is volatile compounds which

were observed in September sample and in subsequent samplings. The volatile compounds

were benzene, toluene, ethylbenzene and xylene (BTEX). Such compounds can pass

through membrane distillation system due to their volatility and pose hazardous to the

operators. Nonetheless, both volatile compound and iron issues can be solved by adding

aeration for pre-treatment unit. Iron would potentially oxidize to precipitate during aeration

105

and could be removed by microfiltration. In the same way, BTEX vaporization is accelerated

when it is aerated.

Based on the calculatations, total membrane area (47 m2), 28 membrane modules (at pilot

scale) were selected in which 25 modules would be operated and 3 modules would be in

standby mode for safety. The arrangement of 28 membrane modules was proposed as shown

in Figure 4.52. CAD files and related dimensions of the complete system are provided

separately in Appendix B .

Figure 4.52 Membrane distillation modules arrangement and pipeline levels

The average flux and energy ratio which were used to design the system were 1.6 kg/m2.h

and 0.624 kW/kg respectively. The wastewater, before using in concentrating step by MD

system, would be passed through the pre-treatment system. At this stage, suspended solids,

oil and grease, and BTEX would be removed.

At the membrane distilling step, wastewater would firstly be heated up until 700C by hot

steam throughout a heat exchanger. A circulative pump would pump the hot feed solution

from the feed tank to heat exchanger and MD modules. Once the feed temperature reached

70 celsius degree, the feed pump would transfer the feed solution to MD module with the

flow rate of 8.6 L per min per module. At permeate side of membrane, the sweeping gas

would be provided by air compressors with the gas velocity inside the fiber of 8.07 m/s. The

permeate flow would be discharged directly into the ambient environment. This system

would be operated by cross-flow mode with the outside-in direction of vapor flux. A

demonstration of designed full scale SGMD plant is shown in Figure 4.53.

T Thermocouple

P Pressure Gauge

Pump

Valve

Flow meter

Air filter/water trap

Heat

exchanger Steam

Air compressor

Feed Tank

1

2 3

1

2

3

Feed pipe level 1

Feed pipe level 2

Feed pipe level 3

a

b

c a Gas pipe level 1

b

b Gas pipe level 2

c Gas pipe level 3

106

Figure 4.53 Sketch Diagram of Full Scale SGMD Plant

107

4.5.2 Financial analysis of full scale SGMD plant

Both investment cost and overall treatment cost of MD plant are expressed in Table 4.20. The

estimated budget information and economic analysis of full scale plant design was provided in

Appendix B.

Table 4.20 Summary of Financial Analysis for Overall System with ±30 % Variation

Components

Treatment cost

(THB)

Estimated Investment Cost

(THB)

Per year Per day Pre-treatment

system MD system

Amortization

cost 521,018 1,427

600,000 5,893,000

O &M costs 104,204 286,0

Membrane

replacement costs 716,800 1,964

Electricity cost 491,962 1,349

Labour cost 164,250 450,0

Total 1,998,234 6,083 6,493,000

*Remark:

- Assume: Plant availability is 90% (the plant operates 328 days per year)

- Treatment cost based on electricity cost, labour cost, O&M cost: 2,291 baht/m3

- In O&M cost, cleaning chemical cost was 32,400 baht per year

Due to the high price of hydrophobic MD module (about 128,000 baht/module), the investment

cost of MD system accounted for higher than 90% total investment cost. Currently, PTT is

treating TEG wastewater by incineration with the treatment cost of 4,500 baht/m3 (135,000

baht/month). Treating by MD system, the calculated treatment cost if basing only on electricity

cost, O&M cost and labour cost is 2,291 baht/m3. By concentrating from 1 m3 to 0.2 m3, the

generating wastewater per month would be 6 m3. This 6 m3 has the TEG concentration of 50%.

It can be further treated by incineration or conducting some more treatment step to reuse TEG

in dehydration process. Overall, the estimated monthly saving when using MD system prior

incineration would be 39,270 baht per month. Although the saving budget was not so high

(30% of current payment), but it would lead to open a potential to recover TEG from

wastewater since MD could be able to concentrate the real wastewater till 99%.

108

Chapter 5

Conclusions and Recommendations

This study mainly focused on optimizing the operational condition of membrane distillation

process for concentrating TEG from wastewater. Two scale of hollow fiber MD were

investigated (0.25 m2 and 2.0 m2). Each membrane module was tested with three solutions:

pure water, synthetic TEG and real wastewater alternately. The best possible condition found

from bench scale study was applied in pilot scale study. Energy consumption analysis in all

experiments was evaluated in term of the ratio of energy consumption/permeate flow. Fouling

analysis was conducted to evaluate the quality of cleaning process. All experimental results

were considered to design a full scale SGMD plant with the capacity of 1 m3/day.

5.1 Conclusions

5.1.1 Pre-treatment system

Pre-treatment unit was designed to remove suspended solids and oil and grease which may

cause operational problem such as membrane fouling and damaged membrane. Suspended

solids was removed by microfiltration (MF). Oil and grease was removed by ultrafiltration The

overall removal efficiency of the pre-treatment unit is 95% and 66% for suspended solids in

term of turbidity and oil and grease respectively. Therefore, the pre-treatment unit concluded

to efficiently remove suspended solids and oil and grease which were the targeted impurities in

real TEG wastewater.

5.1.2 Bench scale hollow fiber membrane distillation study

5.1.1.1 System verification

The optimum condition in terms of energy consumption/Qp and permeate flux were achieved

at feed flow rate of 2.4 (L/min), feed temperature of 70oC, and sweeping gas velocity of 4.7 m/s

(gas inlet flow rate of 0.255 L/min.fiber). At this condition, pure water flux achieved was 3.14

kg/m2.h, the ratio of energy consumption/Qp was 1.09 kWh/kg. The bench scale MD module

has ability to reject 99.99 % of non-volatile substance.

5.1.2.2 Synthetic TEG investigation

The hollow fiber MD was tested at 10% initial TEG concentration. During concentrating TEG

from 10 to 45 %, the permeate flux was in the range from 2.61 to 2.1 kg/m2.h with the ratio of

energy consumption/Qp was approximately 1.29 kW/kg.

The boundary layer resistance of SGMD was not considered as a serious problem. It contributes

only about 10-20% of total resistance and can be removed by rising the system with DI water.

Bench scale MD module has ability to concentrate synthetic TEG till 90%

109

5.1.2.3 Real wastewater investigation

The permeate flux of real wastewater investigation when concentrating TEG from 9.69 to 50%

were in the range from 2.4 to 1.6 kg/m2.h with the ratio of energy consumption/Qp was 1.4

kWh/kg.It could concentrate TEG in real wastewater till 98.01%.

In the total resistance, membrane resistance, boundary layer resistance, fouling resistance

contributed 69.2 %, 7.6%, 23.2 % respectively. The irreversible resistance accounted for 2%.

5.1.3 Pilot scale hollow fiber membrane distillation study

5.1.3.1 System verification

The optimum condition was decided at the gas velocity of 8.07 m/s (gas inlet flow rate of 0.44

L/min.fiber), feed inlet temperature of 70oC. Feed flow rate and sweeping gas temperature have

less effect. The flux achieved was 1.94 kg/m2.h. At this condition, the ratio of energy

consumption/Qp was 0.51 kW/kg.

Synthetic TEG 10% was used as the feed solution to conduct the rejection test. The pilot scale

MD module could reject 99.99 % of non-volatile material.

5.1.3.2 Synthetic TEG investigation

Synthetic TEG (10%) was used as the feed solution. The permeate flux was stable around 1.8

kg/m2.h when the TEG concentration was less than 30 %. When the TEG concentration was in

the range from 30% to 50%, permeate flux was about 1.6 kg/m2.h (average). At higher TEG

concentration (>50%), the flux reduced to less than 1.2 kg/m2.h.

When TEG concentration was in the range from 10 to 60%, the ratio of energy consumption/Qp

was less than 0.7 kW/kg.

Pilot scale MD module has ability to concentrate synthetic TEG till 98.7%

During increasing of TEG concentration (from 10 to 80%), the boundary layer resistance

increased from 51.8 to 133 x 105 m/s. Its contribution to the total resistance increased from 10.4

% to 22.5% which was the main reason of reduction of permeate flux.

5.1.3.3 Real wastewater investigation

The permeate flux was higher than 1.7 kg/m2.h with the TEG concentration less than 15 %. The

flux reduced to the range between 1.4 and 1.65 kg/m2.h at higher TEG concentration from the

hour of 5th and 28th (TEG concentration less than 50%). It reached 0.9 kg/m2.h at the hour of

40th (final TEG concentration was about 69.4 %)

Pilot scale SGMD module has ability to concentrate TEG in real wastewater until 99.1%

110

The contributions of membrane resistance, boundary layer resistance and fouling resistance into

the total resistance were 60.2 %, 10.4 % and 29.4% respectively. After cleaning with chemical

agents, irreversible fouling resistance remained 1.3% total resistance.

To concentrate real wastewater from 10 to 45 % TEG concentration, the required energy ratio

was 0.62 kW/kg. The energy ratio in real wastewater and synthetic TEG were similar at 0.95

kW/kg when concentrating TEG continuously up to 65%.

5.2 Recommendations for Future Study

Membrane distillation is a new technology. Commercialization has not been applied for MD

technology yet. Thus, some recommendations for further research are proposed as followings:

1. The liquid entry pressure (LEP) should be put more intention when using hydrophobic

membrane distillation. This factor helps to ensure that the system is working well with

membrane distillation process. Especially, the relationship between membrane (material,

pore size, thickness, hydrophobicity) and operational condition (feed flow rate, sweeping

gas velocity, temperature, concentration) should be studied deeply.

2. The wastewater discharged from textile industry is becoming a big environmental issue.

Hot textile wastewater is fully infested by dye materials. In textile industry, dyeing step

creates the highest risk for environment (organic dyes, additives and salts). Dyes almost

are non-volatile compounds (mono- and poly- azo). Since the principle of membrane

distillation process is based on the different in vapor pressure between feed side and

permeate side, it is a great potential to recover valuable materials from hot dyeing solution.

Further study should focus on practicalization of recovering dye compounds (i.e.

methylene blue-C16H18ClN3S.3H2O) from synthetic dyeing wastewater. Membrane

distillation can be used to concentrate dyes according to the temperature of dye bath

solution (usually 80-90oC) ,

3. Membrane distillation hybrid system: since it is not necessary to heat the feed solution till

boiling point of its material, the combination between membrane distillation system

(especially heating step) and solar energy system should be studied in depth. According

to the state of art of membrane distillation technology, energy consumption has become

a vital factor that make it become comparable to other conventional process (i.e. RO, MF,

UF). However, there is potential to reduce energy consumption by using solar energy (free

natural energy source).

4. The best available cleaning method for hydrophobic membrane distillation has not been

studied sufficiently. This includes chemical agents, concentration and procedure. It is a

demand for this technology on protect the hydrophobicity of membrane. Further studies

should clarify the interaction between chemical agents and hydrophobic characteristic.

The changes of water contact angle on membrane surface after conducting cleaning step

should be clarified.

111

5. Extensive study should be conducted to understand the effect of the fiber length in SGMD

process. The shorter hollow fiber might lead to better MD performance (in terms of both

concentrating efficiency and energy consumption). Moreover, the effect of hollow fiber’s

thickness should also be studied sufficiently. The membrane thickness has significant

effect on heat loss during membrane distillation process.

112

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Appendix A

Experimental Materials and Activities

118

Figure A.1 Pre-treatment system

Figure A.2 Bench scale SGMD system

Feed

tank

Electric

heater Feed

pump

Control

box

MD

module

1. Raw WW Tank

2. MF Module

3. Intermediate Tank

4. UF Module

5. Pre-treated WW

6. Control box

119

Figure A.3 Bench scale membrane distillation module (0.255 m2)

Figure A.4 Air and liquid flow rate measurement

Air flow

meter

Liquid flow

meter

120

Figure A.5 Pilot scale SGMD system

Figure A.6 Pilot scale MD module (2 m2) and acrylic feed ruler tube

1. Heater

2. Heat exchanger

3. Feed tank

4. Membrane module

5. Sweeping gas inlet

6. Control box

Pilot MD

module Ruler tube

Feed tank

121

Figure A.7 Energy consumption measurement

Figure A.8 Water condensed at the outlet of membrane module

122

Figure A.9 Student operating and maintaining the system

123

Appendix B

Full Scale Sweeping Gas Membrane Distillation Plant Design

Calculations and Drawings

124

Based on the results of this study, a full scale SGMD system was designed for application at

the Gas Separation Plant. The operational capacity of this system is 1 m3/d. All the costs

(investment, operation and maintenance) was estimated based on the market price at the

designing stage. Operational conditions was the optimum conditions from this study.

1 Design Criteria For Full Scale Pre-treatment Plant

Predominantly, the pre-treatment unit is designed to remove suspended solids, oil and grease

which may cause operational problem such as membrane fouling and damaged membrane. The

basic parameter was analyzed in these three samples; raw wastewater, after microfiltration and

after ultrafiltration. The experience gained at pilot scale a schematic (Figure A.1) and design

calculations/considerations are presented below.

Figure B.1 Schematic diagram of pre-treatment unit

1.1 Storage Tanks

1.1.1 Raw Wastewater Tank

Initially the raw wastewater coming from the process needs to be stored in tank. This tank is

designed as a retention/ equalization tank. As per information provided by PTT. TEG

wastewater is 1 m3/day with a concentration of TEG varying from 8 – 20%. But in the current

case the TEG concentration would be assumed to be 10% in the wastewater.

Thus the wastewater influent flow rate: Q = 1m3/d

As a safety factor, assume a storage/ hydraulic retention time: HRT = 3 days

Raw wastewater

tank

D

Intermediate tank Pre-treated

tank

UF Module MF Module

P

Air

P P


Pump Valve

125

Thus the required volume of the tank would be (V) =5 m3

The readily available product with this requirement is presented below with the image as figure

B.3

Choose

Company: Mixer Direct, Inc.

Dimensions (mm): a = 1000, b = 200, c = 4000 , d = 200 , e = 420, f = 63.5, g = 40, h = 40

Inside diameter: 960 mm

Thickness: 20 mm

Nominal capacity: 3250 L

Material: Stainless steel

Figure B.2 TEG Wastewater Tank (custom design)

1.1.2 Intermediate Tank

126

The purpose of this tank is just to act as an intermediator tank before UF application thus a big

size tank is not necessary. This tank was controlled by level sensors the details can be found

below.

Required volume of Tank: V = 0.5 m3

Choose


Dimensions (mm): a = 650, b = 200, c = 1500, d = 200, e = 420, f = 63.5, g = 40, h = 40,


Thickness: 20 mm



Maximum Level sensor: 400 mm (from the top)

Minimum Level sensor: 420 mm (from the bottom)

Operating volume: 430 L

Figure B.3 Appearance of TEG Wastewater Tanks (custom design)

127

1.1.3 Pre-treated Tank

Required volume of Tank: V = 1.3 m3

Choose

Compagnie: Mixer Direct, Inc.

Dimensions (mm): a = 1000, b = 200, c = 1600, d = 200 , e = 420, f = 63.5,

g = 40, h = 40,


Thickness: 20 mm






1.2 Microfiltration and ultrafiltration module design

The overall design of the pre-treatment system would be 1 m3/d. Thus a specific membrane

module are provided based on these results.

1.2.1 MF module

The flux is determined by the properties of membrane, driving force and characteristic of the

feed.

Assume: Water recovery of MF/UF are 99% respectively

→ Permeate flow from MF is 1.01 m3/day.

Permeate flux = 150 L/m2.h, Filtration duration 8 hours per day

Permeate flow: Qp = 1.01 m3/ 8 h = 126.3 L/h

Permeate flux: 50 L/m2.h

→ Total membrane area required: A = Qp / J = 126.3 (L/h) /50 (L/m2.h) = 2.5 m2

Choose:

Company: Microdyn-Nadir (Xiamen) Co., Ltd, China

Model: MD 070 FP 2L

Membrane material: Polypropylene-Hollow fiber modules

Filtration surface Area: 2.2 m2

Membrane pore size: 0.2 µm

Hydrophobicity: Hydrophilic

Shell material: Polypropylene

Inner diameter: 0.6 mm

Module length: 1400 mm

Number of module: 2

128

1.2.2 UF module

Assume: The system was operated 8 h/day

Permeate flow: Qp = 1 m3/ 8 h = 125 L/h

Permeate flux: 40 L/m2.h

→ Total membrane area required: A = Qp / J = 125 (L/h) /40 (L/m2.h) = 3.125 m2

Choose

Company: Pall Corporation,

Model: 60P37-30

Membrane material: Ceramic

Porosity: >30%

Length: 1020 mm

Number of elements: 19

Filtration surface Area: 4.6 m2

Membrane pore size: 50 nm

Channel size: 4 mm

Hydrophobicity: Hydrophilic

Shell material: L316 Stainless steel

Gaskets material: PTFE

1.3 Pre-treatment Process Operation

1.3.1 Aeration of BTEX Removal

As BTEX concentration in the real TEG wastewater, each of the compound’s concentrations

are listed in Table B.1.

Table B.1 BTEX Concentration in 10% TEG Real Wastewater and Henry’s Law

Constant

BTEX Compound Concentration

(mg/L)

Henry’s Law Constanta at 25oC

(atm)

Benzene 26.6 306

Toluene 14.1 372

Ethylbenzene ND 486

m,p-Xylene 7.4 382b

o-Xylene 14.7 293

*Remark: ND – Not detect, a – source: Freeman (1990)

Stripping factor can be calculated by Equation below (Freeman, 1995)

129

SF = (H

Pt)(

G

L)

Where:

H = Henry’s constant (atm)

Pt = Total system pressure

G = Gas flow rate (mol/h)

L = liquid flow rate (mol/h).

Stripping factor (SF) should be in the range from 5 to 20 (Freeman, 1995). Once the removal is

completed, SF equals to 1. The higher Henry’s constant, the lower solubility in water of that

compound. Thus, o-Xylene is selected to calculate SF.

From ideal gas laws, density of air can be calculated as following:

PV = nRT →n

V= ρa =

P

RT

At 25oC, 1 atm.

ρa =1 atm

8.20578 × 10−2 L. atmK. mol

× 298.15 K= 0.041

mol

L

Density of water:

ρw = 1000 g

L×

mol

18 g= 55.56

mol

L

From Equation 1, choose SF =5:

5 = (293 atm

1 atm) × (

G

L)

→ (G

L) = 5 × (

1atm

293 atm) ×

mol h⁄

mol h⁄×

55.56 mol L⁄

0.041 mol L⁄= 23.13 (

L h⁄

L h⁄)

Thus, At SF = 5, the air-to-water ratio is 23.13

Aeration was applied directly at the raw wastewater tank via an air diffuser.

1.3.2 Filtration Step

130

Table B.2 Pre-treatment Process Operating Conditions

Operating conditions Unit Value

MF UF

Maximum feed pressure Bar 10 6.25

Maximum operating TMP Bar 1 2.1

Filtrate flux (at 25oC) L/m2.h 50 40-120

Temperature oC < 60 < 40

Operating pH range 0 -14 2-11

A normal operation process includes two steps: filtration and backwash. The range of operating

cycle is from 20 to 60 minute. During removing contaminants period, the transmembrane

pressure will rise. Thus, a back wash step is necessary to recover membrane permeability.

2 Design Considerations for Full Scale Sweeping Gas Membrane Distillation Unit

2.1 Storage Tanks

The capacity of full scale MD system: 1 m3/day


Dimensions (custom design, mm): a = 1000, b = 200, c = 1500, d = 200 , e = 420, f = 63.5,

g = 40, h = 40,


Thickness: 20 mm






2.2 Hollow fiber module

Qo = 1 m3/d; Co = 10%

Assume the final TEG concentration is C1 =50% (theoretical value) → Q1 = 0.2 m3/d

→ Permeate flow = Qo - Q1

= 1-0.2

= 0.8 m3/d

= 80 L/h (the system operates 10h/day)

131

Permeate flux: 1.2 – 1.9 kg/m2.h (when concentrating real wastewater from 10 to 50% TEG

concentration)

Total membrane area required:

A = 80 (L h⁄ )

1.6 (kg m2. h)⁄× 1 (kg L) = 50 m2⁄

Table B.3 Specification Hollow Fiber Membrane Distillation


Company Name Sumitomo Electric Industries, Ltd.

Membrane Name PM-X215

Type No. 13001-2

Module Configuration Hollow Fiber

Membrane Material Polytetrafluoethylene (PTFE)

Type of Membrane Hydrophobic Microporous

Contact Angle 112o

Nominal Pore Size (μm) 0.45

Outside Diameter (mm) 2.03

Inside Diameter (mm) 1.07

Total Length (mm) 1105

Effective Length (mm) 1008

Thickness (μm) 480

Number of Elements 306

Membrane Effective Area (m2) 2

Operating Temperature Range (oC) -100 to 260

Total module needed:

𝑛 =50 𝑚2

2= 25

→ Choose 28 membrane modules (25 working, 3 standby)

132

Table B.4 Membrane Module Specification

Description Characteristics

Type of Membrane Module Hollow Fiber

Module Configuration SGMD

Frame Material Polysulfone

Driving Force Thermal Driven

Inner Space (cm3) 7,300

Pipe Diameter (mm) 25.4 (feed)/6 (permeate)

Number of fiber 306 fibers/module

Dimension of Module (cm)

- Diameter

- Length

9.6

110.5

Operating Temperature Range

(oC) -100 to 149

pH Range 2-13

Membrane System Fabrication

Number of module: 28

CAD files and related dimensions of the complete system are provided separately.

2.3 Gas compressor

Pilot scale module: 306 fiber/module

Optimum Inlet Gas Flow rate: 133 L/min

Gas flow rate per fiber =133 L min⁄

306 fibers= 0.44 L/min

The sweeping gas flow rate needed for pilot scale module (306 fibers/module):

Qgas = 0.44 L/ min × 306 fibers × 25 modules = 3366 L/min

= 3.23 m3/min

Number of gas compressors: 3 (2 operate and 1 standby). These calculations are based on the

optimization results obtained from pilot scale testing.

133

Table B.5 Characteristics of Gas Compressor


Company Name Hitachi

Model OBB – 7.5HB6

Type Bebicon – Oil free

Full load power 7.5 kWh

Motor efficiency 95%

Weight 261 kg

Maximum pressure 0.5 MPa

Free air delivery 2500 L/min

Maximum discharge pressure 1.37 MPa

External dimensions (WxDxH) 1938 x 608 x 1114

Number of compressor 3

Air tank volume 280 L

Operational Mode Pressure switch off

Figure B.4 Sketch Diagram of Full Scale MD Plan

T Thermocouple

P Pressure

Gauge

Pump

Valve

Flow meter

Air filter/water trap

Heat

exchanger

Steam

Air compressor

Feed Tank

1

2 3

1

2

3

Feed pipe level 1

Feed pipe level 2

Feed pipe level 3

a

b

c a Gas pipe level 1

b

b Gas pipe level 2

c Gas pipe level 3

134

2.4 Heat exchange system

2.4.1 Steam calculation

Continuous Heating Processes

In heat exchangers the product or fluid flow is continuously heated.

The mean heat transfer can be expressed as:

q = Cp × ∆T × (m t⁄ )

Where

q = mean heat transfer rate (kJ/s))

m t ⁄ = mass flow rate of the product (kg/s)

cp = specific heat capacity of the product (kJ/kg.oC)

∆T = change in temperature of the fluid (oC)

Assume:

Triethylene glycol 30%,

Specific heat capacity at 70oC: CP = 3.94 kJ/kgoC

∆T = 70oC – 25oC = 45oC

Inner diameter of module: 96 mm

Cross section area created by membrane module:

𝐴𝑚 = 𝜋 ×0.0962

4= 7.24 × 10−3 (𝑚2)

Cross section area created by membrane fibers:

𝐴𝑓 = 𝜋 ×(2.03 × 10−3)2

4× 306 = 9.9 × 10−4 (𝑚2)

Total cross section area for feed solution:

∆𝐴 = 𝐴𝑚 − 𝐴𝑓 = 7.24 × 10−3 − 9.9 × 10−4 = 6.25 × 10−3 (𝑚2)

In sweeping gas membrane distillation configuration, the effect of feed flow rate is negligible.

Feed velocity used in bench scale experimental study: 0.023 m/s

Thus, feed flow rate for pilot scale per module is:

𝑄 = ∆𝐴 (𝑚2) × 𝑣 (𝑚 𝑠)⁄ = (6.25 × 10−3) × 0.023 = 1.4375 × 10−4 (𝑚3 𝑠⁄ ) = 8.6 (𝐿 𝑚⁄ )

Feed flow rate for pilot scale per module = 8.6 L/min

135

When 25 modules work:

𝑄 =8.6 𝐿 𝑚𝑖𝑛⁄ × 25 𝑚𝑜𝑑𝑢𝑙𝑒𝑠 × 1.018 𝑘𝑔 𝐿⁄

60 𝑠 𝑚𝑖𝑛⁄= 3.65 (𝑘𝑔 𝑠)⁄

→ Mean heat transfer rate

q = 3.94 × 45 × 3.65

= 646.76 (𝑘𝐽 𝑠⁄ )

≈ 2,208,379.1 (𝐵𝑇𝑈 ℎ)⁄

2.4.2 Steam requirement

The amount of steam can be calculated:

ms = q / he

Where:

ms = mass of steam (kg/s)

q = calculated heat transfer (kJ/s)

he = evaporation energy of the steam (kJ/kg)

Assume: The pressure of steam is 8 bar → he = 2030 kJ/kg

ms =671.97 kJ/s

2030 kJ/kg= 0.305 kg/s = 1096 kg/h

High temperature source is stream → Choose tube heat exchanger with counter-flow mode

Heat transmission coefficients: Steam to Copper to Water: 205

Steam temperature: 657.824oF (at 8 bars)

Liquid inlet temperature: 77oF (25oC)

Outlet temperature: 160oF (70 oC)

Required surface area:

Q = Heat transfer coefficient × Area × (Steam temp − Outlet temp)

→ Area =2,294,388 BTU h⁄

205 × (657.8 − 160)

= 22.48 ft2 ≈ 2.1 m2

136

Choose:

U-tube heat exchanger, steam to water

Company: Needs custom manufacturer

Diameter: 152.4 mm

No. of Passes: 4

Heat surface: 2.1 m2

Tube side pressure: 10.3 bars

Shell side pressure: 10.3 bars


Note: Even though brass is an excellent material for heat exchangers but in suitability of brass

heat exchanger for wastewater concentration and separation might result in TEG or other

compounds reacting to the heat exchanger material itself.

2.4.3Alternatively electric heater requirement

In Continuous Heating Processes, the heat exchangers the product or fluid flow is continuously

heated.

The mean heat transfer can be expressed as

q = Cp × ∆T × (m t⁄ )

Where

q = mean heat transfer rate (kJ/s))

m t ⁄ = mass flow rate of the product (kg/s)

cp = specific heat capacity of the product (kJ/kg.oC)

∆T = change in temperature of the fluid (oC)

Inner diameter of module: 96 mm

Cross section area created by membrane module:

𝐴𝑚 = 𝜋 ×0.0962

4= 7.24 × 10−3 (𝑚2)

Cross section area created by membrane fibers:

𝐴𝑓 = 𝜋 ×(2.03 × 10−3)2

4× 306 = 9.9 × 10−4 (𝑚2)

Total cross section area for feed solution:

∆𝐴 = 𝐴𝑚 − 𝐴𝑓 = 7.24 × 10−3 − 9.9 × 10−4 = 6.25 × 10−3 (𝑚2)

137

In sweeping gas membrane distillation configuration, the effect of feed flow rate is negligible.

Feed flow rate for pilot scale per module is:

𝑄 = ∆𝐴 (𝑚2) × 𝑣 (𝑚 𝑠)⁄ = (6.25 × 10−3) × 0.023 = 1.4375 × 10−4 (𝑚3 𝑠⁄ ) = 8.6 (𝐿 𝑚⁄ )

Feed flow rate for pilot scale per module = 8.6 L/min

When 26 module works:

𝑄 =8.6 𝐿 𝑚𝑖𝑛⁄ × 26 𝑚𝑜𝑑𝑢𝑙𝑒𝑠 × 1.018 𝑘𝑔 𝐿⁄

60 𝑠 𝑚𝑖𝑛⁄= 3.79 𝑘𝑔 𝑠⁄

Assume:

Triethylene glycol 30%,

Specific heat capacity at 70oC: CP = 3.94 kJ/kgoC

∆T = 70oC – 25oC = 45oC

→ Mean heat transfer rate

q = 3.94 × 45 × 3.79 = 671.97 𝑘𝐽 𝑠⁄ ≈ 2,294,388 𝐵𝑇𝑈 ℎ⁄

Conversion: 1 kWh = 3412 BTU

Thus, the total heat energy needed is:

𝑄𝑒 =2,294,388 𝐵𝑇𝑈 ℎ⁄

3412 𝐵𝑇𝑈= 672.4 𝑘𝑊

3 Pump and pipeline system

3.1 Pre-treatment system

3.1.1 Pump

MF module

- Flow rate: 1010 L/h

- Head capacity: 5 bars = 50.9 meters of water

UF Module

- Flow rate: 1000 L/h

- Head capacity: 6.25 bars = 63.7 meters of water

Pump capacity

138

P = Q. ρ. g. H

3.6 × 106. η=

1.01 m3 h × 1125 kg m3 × 9.81 m s2⁄ × 63.7 m⁄⁄

3.6 × 106 × 0.65

= 0.3 KW = 0.4 Hp

Where:

ρ: Specific density of TEG wastewater, choose ρ = 1125 kg/m3 (for pure TEG)

η: pumping coefficient, choose η=0.65

Table B.6 Characteristics of Pre-treatment Pump


Company Name ProMinent, Germany

Model DFCa 040

Pump capacity 1583 L/h

Pressure Max. 8 bars

Connector size 1.5 inch

Weight 89 kg

Feed rates 0.43- 15.83 L/rev

Self-priming up to 8m

Hose diameter (inside) 35 mm

Number of Pump 2

3.1.2 Pipeline system

Flow rate: 1000-1010 L/h

Wastewater velocity: v = 1-5 m/s (Metcalf and Eddy, 2005)

Assume v= 2 m/s

D = √4. Q

π. v= √

4 × 1010 L h⁄

π × 2 m s⁄= 13.36 mm

Choose:

Company name: MISUMI Thailand

Material: Vinyl Chloride

Model: PVCH No. 13

139

Outer diameter: 18 ± 0.2 mm

Thickness: 2.5 ± 0.2 mm

Inner diameter (Reference value): 13 mm

Re-check the feed velocity:

v = 4. Q

π. D2=

4 × 1010 L h⁄

π × 0.0132= 2.1 m s⁄ (satisfied)

3.2 Membrane Distillation Module

3.2.1 Pump

Feed flow rate per module: 8.6 L/min = 516 L/h

Total feed flow rate for whole system: Q = 516 L/h x 25 modules =12900 L/h = 12.9 m3 h⁄

MD operating pressure: 1 atm

Pump capacity

P = Q. ρ. g. H

3.6 × 106. η

= 12.9 m3 h × 1125 kg m3 × 9.81 m s2⁄ × 10m⁄⁄

3.6 × 106 × 0.65

= 0.61 kW

= 0.81 Hp

Table B.7 Characteristics of Pre-treatment Pump


Company Name ProMinent, Germany

Model DFDa 100

Pump capacity 15 000 L/h

Pressure Max. 15 bars

Weight 1100 kg

Feed rates 0.3-20 L/rev

Self-priming up to 8m

Hose diameter (inside) 100 mm

Number of Pump 1

3.2.2 Pipeline system

Feed characteristic: High temperature (70oC), low pH (<5)

140

→ Choose Stainless steel 316 piping

i. TEG wastewater feed pipeline

- The main feed pipe at 1st level: from Pump to 4 pipe line at 2nd level

Flow rate: 12,900 L/h

Wastewater velocity: v = 1-5 m/s (Metcalf and Eddy, 2005). Assume v= 2 m/s

D = √4. Q

π. v= √

4 × 12,900 L h⁄

π × 2 m s⁄= 47.7 mm

Choose:

Material: ASTM A316 | Sweglok SS 316

Outer diameter: 60.03mm

Thickness: 2.77 mm


Weight: 3.93 kg/m


v = 4. Q

π. D2=

4 × 12,900 L h⁄

π × 0.052= 1.8 m s⁄ (satisfied)

- The 2nd level pipe (from 1st level pipe to the 3rd level pipe - connect to the module)

Flow rate Q = 12,900 (L/h) / 4 = 3225 L/h


D = √4. Q

π. v= √

4 × 3225 L h⁄

π × 2 m s⁄= 23.88 mm

Choose:

Stainless pipe

Material: ASTM A316

Outer diameter: 33.4 mm

Thickness: 2.77 mm


Weight: 1.28 kg/m


141

v = 4. Q

π. D2=

4 × 3225 L h⁄

π × 0.0252= 1.8 m s⁄ (satisfied)

- The 3rd level pipe (from the 2nd level pipe to each membrane module)

Flow rate Q = 8.6 L/min


D = √4. Q

π. v= √

4 × 8.6 L min⁄

π × 2 m s⁄= 9.55 mm

Choose:

Stainless pipe

Material: ASTM A316


Thickness: 1.65 mm

Iner diameter (Reference value): 10 mm

Weight: 0.85 kg/m


v = 4. Q

π. D2=

4 × 8.6 L min⁄

π × 0.012= 1.8 m s⁄ (satisfied)

ii. Sweeping gas pipeline

Sweeping gas characteristic: Natural atmospheric gas

- The main gas pipe at 1st level: from Pump to 4 pipe line at 2nd level

Flow rate: 2028.8 L/min

Gas velocity: v = 10-20 m/s (Metcalf and Eddy, 2005). Assume v= 15 m/s

D = √4. Q

π. v= √

4 × 2028.8 L min⁄

π × 15 m s⁄= 53.57 mm

Choose:

Stainless pipe

142

Material: ASTM A316

Outer diameter: 60.03mm

Thickness: 2.77 mm


Weight: 3.93 kg/m


v = 4. Q

π. D2=

4 × 2028.8 L min⁄

π × 0.052= 17.22 m s⁄ (satisfied)

- The 2nd level pipe (from 1st level pipe to the 3rd level pipe - connect to the module)

Flow rate Q = 2028.8 (L/min) / 2 = 1014.4 L/min


D = √4. Q

π. v= √

4 × 1014.4 L min⁄

π × 15 m s⁄= 37.8 mm

Choose:

Stainless pipe

Material: ASTM A316


Thickness: 2.77 mm


Weight: 3.11 kg/m


v = 4. Q

π. D2=

4 × 1014.4 L min⁄

π × 0.042= 13.45 m s⁄ (satisfied)

- The 3rd level pipe (from the 2nd level pipe to each membrane module)

Flow rate Qgas = 0.255 L/ min × 306 fibers = 78.03 L/min


D = √4. Q

π. v= √

4 × 78.03 L/min

π × 15 m s⁄= 10.05 mm

Choose:

Stainless pipe

143

Material: ASTM A316


Thickness: 1.65 mm


Weight: 0.85 kg/m


v = 4. Q

π. D2=

4 × 78.03 L/min

π × 0.012= 16.6 m s⁄ (satisfied)

4 Clean in Place (CIP)

4.1 Pre-treatment system

After a certain operating time, the MF/UF membrane system needs to be cleaned. This is on

demand of membrane filtration.

Top Backwash Step

When the TMP exceeds 1 bar above starting TMP (at same temperature), the membrane system

should be stop to conduct the cleaning step. The clean water is provided from the inside space

of hollow fiber to the outside. The wastewater is drained out from the top of module housing.

This step is carried out to remove particulates at the highest concentration.

Duration: 15 minutes

Cleaning solution source: Tap water

Backwash pressure: 8 bars

Amount of tap water required

a) MF module

- Back wash flux: 50 L/m2.h

- Volume of water required:

V𝑀𝐹 = 50 (L m2. h)⁄ × 21 m2 × 0.25 h

= 262.5 L

b) UF module

- Back wash flux: 40 L/m2.h

- Volume of water required:

V𝑈𝐹 = 40 (L m2. h)⁄ × 33 m2 × 0.25 h = 330 L

Bottom Backwash Step

144

Similar with top backwash, the clean water flow pass membrane pores by inside-out mode. In

this step, backwash water is drained off from the bottom of module. The purpose is to remove

heavier materials which cannot remove by top backwash.

Forward flush

Depending on the degree of fouling, backwash step can be repeated a numerous times (normally

from 3 to 8 times). After that, the pre-treatment system was returned back to the normal

operating mode.

Chemical enhanced cleaning

Duration: 120 minutes (recycle and soak)

Cleaning Solutions: 0.2 % HCl, and 0.1% NaOH

Cleaning flux per Module: 1.0-1.5 m3/h

Temperature range: 10 to 40oC

Frequency: In the range from 1 to 3 months.

4.2 Membrane Distillation System Cleaning

Pure water cleaning

In MD technology, fouling phenomena is not a serious operating problem. Deposited particulars

can be removed by the turbulent velocity of water flow. Thus, tap water is used to clean the

fiber surface and shell surface.

Duration: 30 minutes

Frequency: 1 time per week

Cleaning solution: Pure water

Volume of pure water: 300 L

Cleaning procedure: The cleaning mode is by-pass. Pure water will not be heat up and sweeping

gas will not be provided. The cleaning wastewater was recirculated back to the feed tank. After

that, it was drained out to the sewer system.

This step can be repeated a numerous times to achieve a good cleaned module.

Chemical cleaning

When operating with wastewater (either synthetic or real wastewater), boundary layer was

formed on the membrane surface. Consequently, fouling resistance will contribute to total

resistance. It will result in decreasing of permeate flux along with timeline of operational

process.

Chemical agents which will use to clean the hydrophobic membrane are a mixture of oxalic and

citric acid. This cleaning solution has been proven as an efficient solution to clean the

hydrophobic membrane distillation. Time of cleaning period and chemical concentration are

shown in Table B.8.

145

Table B.8 Chemical Concentration for Membrane Cleaning

Step Cleaning

procedure

Concentration

(mg/L) pH

Cleaning

Time (h) Purpose

1 Water (1) - 6.15 0.5 Recoverable fouling

2 Oxalic/Citric

0.1wt. % Oxalic acid

and

0.8wt. % Citric acid

2.2 6 Reversible fouling

3 Water (2) - 6.15 0.5 Washing the remain

chemical

4 Sunlight 24 Dry-out

The procedure is similar with pure water cleaning step. However, chemicals are weighted and

put directly into the feed tank. The chemical solution was mixed evenly by the feed pump.

Frequency: 1 time per month

Total chemical liquid volume: 300 L

Amount of acid per time of cleaning: 0.3 kg Oxalic acid (0.1 wt%) and 2.4 kg Citric acid (0.8

wt%)

Clean water storage tank

Total minimum volume required for both pre-treatment system and MD system:

V = VMF + VUF + VMD = 262.5 + 330 + 300 = 892.5 L Choose

Model: CEN2K0A-M107

Manufacturer: PAKCO International Co., Ltd.

Material: Polyethylene (PE)

Dimension: B = 395mm, C = 1300 mm, H = 1900 mm, h = 1650 mm, K = 395 mm, O = 90

mm. Thickness = 10 mm.

Capacity: 2 m3

146

Figure B.5 Tap water tank

5 Full scale budget information and economic analysis

5.1 Estimated budget information of full scale plant

Concentrating from 10 to 50% for 1 m3

Vinitial = 1 m3 ;Vfinal = 0.2 m3 → ∆V = 0.8 m3 of water (approximate to 800 kg)

Table B.9 Estimated budget summary for overall system with ±30 % variation

Components Details Price/unit Number Estimated Total

Membrane

module 2 m2 surface area / module 128000 28 3,584,000

Piping Sanitary piping, valves, joints

splitters etc. 500,000

Electrical

connections

Thermocouples, LV

controllers, Control box etc. 300,000

Tanks 1,000,000

Pretreatment

system 600,000

Air

Compressor 1225 L/min 3 400,000

Flow meters Max 100 LPM 6 400,000

Pressure

gauges 1 bar 4000 25 100,000

147

Pressure

gauges 5 Bar 3000 3 9,000

Pumps

100 LPM (Circulation pump

for MF) 1 60,000


for UF, 5 Bar) 1 60,000


for MD) 1 100,000

Heat exchanger Maintain temp 70 Deg C for

0.5 m3 water 1 400,000

Total 6,493,000

5.2 Energy Analysis

In technology innovation, energy consumption of MD process is an important factor for this

energy intensive technology. The components consuming energy in MD process are: thermal

energy to heat the feed liquid, energy required for circulation pump and compressor. The heat

energy requirement is more than 90% of the total energy consumption of MD process (M.

Khayet and Matsuura, 2011)

The balance of heat energy is shown in Equation below

𝑄𝑚 = 𝑄𝑐 + 𝑄𝑣

Where Qm, Qv, Qc are total heat transfer in MD process, heat transfer by mass transfer and heat

transfer by conduction of membrane respectively

𝑄𝑣 = 𝐽𝑤 × ∆𝐻𝑣,𝑤 = 1.95 𝑘𝑔 𝑚2. ℎ⁄ × 2270 𝑘𝐽 𝑘𝑔⁄ = 4426.5 𝑘𝐽 𝑚2. ℎ⁄

𝑄𝑐 =𝑘𝑚

𝛿× (𝑇𝑚𝑓 − 𝑇𝑚𝑝) =

0.14 𝑊 𝑚. 𝐾⁄

0.00048 𝑚× (67.7 − 66.6) = 320.8

Thermal efficiency, η

η =Qv

Qv + Qc× 100 =

4426.5

4426.5 + 320.8× 100 = 93.2 %

𝑄𝑐 is considered as heat loss. The thermal efficiency in SGMD is very high (93.2%) while that

value for DCMD is between 50-80% (Khayet and Matsuura, 2011). Thus, it can be concluded

that the heat loss is not significant in MD process.

148

Figure B.6 Energy consumption/Qp of continuously-fed real wastewater investigation

The average energy required to concentrate TEG in real wastewater from 10 to 60 % is 0.624

kW/kg (624 kW/m3). Energy consumption had been evaluated for DCMD and VMD at lab scale

(Criscuoli et al., 2008). The authors conducted a study on plate-and- frame membrane module

(membrane area of 40 cm2). The lowest ratio of energy consumption/Qp obtained were 3546.3

kW/m3 and 1108.4 kW/m3 in DCMD and VMD respectively. The performance of VMD was

proven to be better than DCMD (based on permeate flux, energy consumption and thermal

efficiency). As yet, almost studies were done theoretical modeling. Thus, it is very important to

conduct study on realistic energy consumption.

5.3 SGMD Costs Evaluations

The annual operating cost is the yearly purchase to own and operate MD system. This amount

include amortization, O&M costs and membrane replacement cost.

5.3.1 Amortization or fixed charges

This annual interest payment is normally charged to cover the initial capital cost when the fund

is borrowed.

Amortization factor (a) is calculated as

a =i × (1 + i)n

(1 + i)n − 1

5

15

25

35

45

55

65

0.0

0.2

0.4

0.6

0.8

1.0

0 5 10 15 20 25 30 35 40 45

TE

G c

on

cen

trat

ion

(%)

En

ergy r

atio

(kW

/kg)

Time (h)


149

Where i, n are the annual interest rate (%) and the year of life time of MD system.

Assume i = 0.05 (5%) and n = 20 years (Banat and Jwaied, 2008)

a =0.05 × (1 + 0.05)20

(1 + 0.05)20 − 1= 0.080243 (year−1)

Therefore, the annual fixed charges is:

Afixed = a × capital cost = 0.080243 × 6,493,000 baht = 521,017.8 baht year⁄

5.3.2 O &M costs

The annual maintenance and operation costs are estimated to account for 20% of the plant

annual payment (Khayet and Matsuura, 2011)

𝐴𝑂&𝑀 = 0.2 × Afixed = 0.2 × 521,017.8 = 104,203.6 baht year⁄

5.3.3 Membrane Replacement costs

This is the cost required for membrane replacement. It is a function of production process.

Membrane replacement cost is calculated as 20% of the membrane module cost.

𝐴𝑀𝑅 = 0.2 × MC = 0.2 × 3,584,000 = 716,800 baht year⁄

5.3.4 Plant Availability (f)

Plant availability means the total working time of the plant per year. It can be assumed to be

90% per year (Khayet and Matsuura, 2011).

5.3.5 Electricity cost (𝑨𝒆𝒍𝒆𝒄𝒕𝒓𝒊𝒄)

A𝑒𝑙𝑒𝑐𝑡𝑟𝑖𝑐 = 𝑐 × 𝜔 × f. M. 365

Where 𝑐, 𝜔, f, M are electric cost (3 baht/kWh), specific energy consumption (kWh/m3), plant

availability (90%) and plant capacity (1 m3/d) respectively.

Specific energy consumption (0.624 kWh/kg), concentrate TEG from 10 to 50%. Thus the total

volume reduce from 1000 L (1 m3) to 200 L

𝜔 = 0.624 × (1000 − 200) = 499.2 kWh/m3

Thus,

150

A𝑒𝑙𝑒𝑐𝑡𝑟𝑖𝑐 = 𝑐 × 𝜔 × f. M. 365

= 3 baht kWh⁄ × 499.2 kWh m3⁄ × 0.9 × 365 days year⁄ × 1 m3 day⁄

= 491,961.6 baht year⁄

5.3.6 Labour cost (Alabour)

A𝑙𝑎𝑏𝑜𝑢𝑟 = 𝑔. f. M. 365

Where g (500 baht/m3) is specific labour cost

A𝑙𝑎𝑏𝑜𝑢𝑟 = 500 baht m3⁄ × 0.9 × 365 days year⁄ × 1 m3 day⁄ = 164,250 baht year⁄

Therefore, the total annual cost (Atotal)

𝐴𝑡𝑜𝑡𝑎𝑙 = 𝐴𝑓𝑖𝑥𝑒𝑑 + 𝐴𝑂&𝑀 + 𝐴𝑀𝑅 + A𝑒𝑙𝑒𝑐𝑡𝑟𝑖𝑐 + A𝑙𝑎𝑏𝑜𝑢𝑟

= 521,017.8 + 104,203.6 + 716,800 + 491,961.6 + 164,250 = 1,998,232.9 baht year⁄

Treatment cost based on the total annual cost

TC =Atotal

f. M. 365=

1,998,232.9 baht year⁄

0.9 × 365 days year⁄ × 1 m3 day⁄ = 6,082.9 baht m3⁄

Treatment cost based on Electricity cost, Labour cost, O &M costs

TC =156,305.3 + 491,961.6 + 104,203.6 baht year⁄

0.9 × 365 days year⁄ × 1 m3 day⁄ = 2290.6 baht m3⁄

Chemical cleaning of MD system

Frequency: 1 time per month

Total chemical liquid volume: 300 L

Amount of acid per time of cleaning: 0.3 kg Oxalic acid (0.1 wt%) and 2.4 kg Citric

acid (0.8 wt%)

Total amount of chemicals using per year:

W𝑎𝑐𝑖𝑑𝑠 = W𝐶𝑖𝑡𝑟𝑖𝑐 + W𝑂𝑥𝑎𝑙𝑖𝑐 = (2.4 𝑘𝑔 + 0.3 𝑘𝑔) × 12 𝑚𝑜𝑛𝑡ℎ𝑠 = 32.4 𝑘𝑔

Chemicals cost per year

𝐴𝑐ℎ𝑒𝑚𝑖𝑐𝑎𝑙𝑠 = 32.4 𝑘𝑔 × 1,000 𝑏𝑎ℎ𝑡 𝑝𝑒𝑟 𝑘𝑔 = 32,400 𝑏𝑎ℎ𝑡

This chemical cost was included in annual O&M cost

151

Appendix C

Operational Manual of Pilot scale Sweeping Gas Membrane Distillation System

152

1. General information

1.1 Semi-pilot scale of pretreatment unit

The pre-treatment unit is a combined membrane system, hollow fiber ultrafiltration

membrane (UF) followed by cartridge microfiltration membrane (MF). This unit played

as a physical filtration barrier, to remove suspended solids (by MF), oil and grease (by UF).

The imagine of pre-treatment unit is expressed in Figure C.1. The air-to-water ratio was set

at 24. The purpose of adding air diffuser is to remove BTEX and transfer iron ion to ferric

hydroxide (Fe(OH)3)

Figure C.1 Schematic diagram of the semi-pilot scale pre-treatment unit

7. Raw WW Tank

8. MF Module

9. Intermediate Tank

10. UF Module

11. Pre-treated WW

12. Control box

153

1.2 MD Semi-pilot Scale System

Mode operation of the study was cross flow mode.

Figure C.2 Semi-pilot Scale Membrane Distillation System

A 60-L stainless steel tank is used as feed tank.

Figure C.3 Feed Tank and ruler tube of the MD System

7. Heater

8. Heat exchanger

9. Feed tank

10. Membrane module

11. Sweeping gas inlet

12. Control box

154

2. System Operation Overview

2.1 Pre-treatment System

Figure C.4 Pre-treatment system

2.1.1 Oparational Steps

1. Check all the tubes and connectors first.

2. Check the MF membrane housing is tight enough.

3. Fill the Waste water to Tank 1. (200 L)

4. Start air stripping- air is applied to the water in the ratio air-to-water of 24.

It need around 5 hours air stripping for 150 L with 13.6 L/m air flow. This depends

on Initial BTEX concentration.

5. Switch on the system. (Control box switch)

6. Keep fully open the baypass valves initially. (Both MF and UF)

7. Switch on the “Run switch”.

8. Then slowly close the bypass valve and see the pressure gauge 1. Pressure should

be about 1 bar. Adjust pressure with bypass valve.

9. Press the air release valve on MF membran housing and release the air initially.

When water come out from the releasin button, un press it.

10. When Pump 2 works, slowly adgust the pressure on UF membranes by UF bypass

valve. Pressure should be below 2-4 bar.

11. Check the permeate line whether fulx is commig or not. Permeate will come after

adjusting the pressure on UF membrane.

12. System will automatically run after this.

2.1.2 Cleaning pre-treatment system membranes

1. Remove the housing of the cartridge MF membrane.

Raw wastewater tank

D

Intermediate tank Pre-treated tank

UF Module MF Module

P

Air

P P


Pump Valve

Tank

1 Tank

2

Tank

3

155

2. Safely remove the ceramic MF membrane.

3. Simply add water and remove solids with a sponge.

4. Sponge it only one direction.

2.2 MD Semi-Pilot Scale System

MF Filter Cartridge after

filtration

MF Filter Cartridge after

cleaning

156

T

L

F/T

T

T

T

Permeate

tank

Heat

exchanger

Heat

exchanger

Heat

exchanger

Chiller

Air compressor

Chiller

Heater

Feed tank

Hollow fiber

module

F/T TPump Pressure gauge ValveThermocoupleAir fi lter/water traper L Level sensor

Permeate

measuring tube

Flow meter

1 2

3

4

Figure C.5 Schematic diagram of the semi-pilot scale MD system

2.2.1 Starting Up

1. Fill the pre-treated wastewater to feed tank.

2. Check the water bath water level initially and maintain the level with DI

water.

3. Switch on the heater and adjust the temperature. Then Switch on the

circulation pump of the heater.

4. Valve 1, 2 and 3 are to be closed and valve 4 is to be opened.

5. Then, feed pump is to be turned on and circulates the solution directly to

heat exchanger

2.2.2 Distilling

1. After reach desired temperature, valve 1 and 4 are need to close and valve

2 and 3 are to be opened.

2. Circulate feed solution through the membrane.

3. Then, ambient air is to be pumped into the membrane.

4. At this stage, inlet pressure of air has to be carefully adjust with pressure

regulator and matched with inlet of feed solution.

2.2.3 Permeate Flux Measurement

157

After distilling process, the level of feed tank is to be initially and hourly noted.

Then permeate flux can be calculated by the following equation:

𝑃𝑒𝑟𝑚𝑒𝑎𝑡𝑒 𝑓𝑙𝑢𝑥 (𝑘𝑔

𝑚2. ℎ) =

𝐿𝑉𝑖 − 𝐿𝑉0

𝑡𝑖 − 𝑡0 × 0.8

Where LV1 and LV0 are water level (cm) at time i and initial time respectively.

2.2.4 Finishing

1. When the experiment is finished, compressed air is the first to be closed.

2. Then, feed pump use to be closed.

3. To discharge the feed solution, valve 1 is to be opened.

2.3 Maintenance of MD System

2.3.1 Membrane Cleaning

MD membrane need to be clean after every batch.

1. Follow table 2.1 to do membrane cleaning.

2. Fill the feed tank with each solution in each steps and run the feed circulation

pump for 20 min.

3. Valve 1 and 4 are to be closed and valve 2 and 3 are to be opened.

4. After each run, feed pump use to be closed.

5. To discharge the feed solution, valve 1 is to be opened.

Table C.2 Steps of Chemical Cleaning

Step Chemicals Concentration

(mg/L)

pH Cleaning

Time (h)

Purpose

1 Water (1) - 6.15 0.5 Recoverable

fouling

2 Oxalic/Citric 0.1wt. % Oxalic

acid

and

0.8wt. % Citric

acid

2.2 6 Reversible

fouling

3 Water (2) - 6.15 0.5 Washing the

remain

chemical

4 Sunlight 24 Dry-out

3 Precautions for Safe Handling and Use

158

3.1 Precautions

Keep away from heat, sources of ignition. Empty containers pose a fire risk, evaporate the

residue under a fume hood. Ground all equipment containing the compound. Do not ingest.

Do not breathe gas/fumes/vapor/spray. Avoid contact with eyes if ingested, seek medical

advice immediately and show the container or the label.

Setup the Inhalation hazard symbol at the pre-treatment system area

Figure C.6 Inhalation Hazard Symbol

3.2 Storage

Keep container dry. Keep in cool place. Ground all equipment containing the compound.

Keep container tightly closed. Keep in a cool, well-ventilated place. The compound should

be stored away from extreme heat and away from strong oxidizing agents.

3.3 Personal protection

3.3.1 Respiratory protection

Half Face piece Respirator (gas mask) with a chin-style

Figure C.7 3M Half Facepiece Respirator 7501

Product information

159

Manufacturer: 3M Occupational Health and Environmental Safety Division (OH &

ESD), USA

Material: Silicone

Size: Small

Filter cartridge: Model 6003-NIOSH, use to avoid certain organic vapors, chlorine,

hydrogen chloride, and sulfur dioxide or hydrogen sulfide or hydrogen fluoride

Figure C.8 3M Half Facepiece Respirator with Cartridge 6003

Using guideline (suggested by the manufacturer)

Donning

(1) Adjust head cradle size as needed to fit comfortably on head.

(2) Place the respirator over the mouth and nose, then pull the head harness over the

crown of the head.

(3) Grasp the bottom straps, place them at the back of the neck and hook them together.

Pull the ends of the straps to adjust the tightness.

(4) Do not over-tighten. Perform a positive and/or negative pressure user seal check

each time the respirator is donned.

User seal check

Always check the seal of the respirator on your face before entering a contaminated

area. Inhale gently.

(1) If you feel facepiece collapse slightly and pull closer to your face with no leaks

between the face and facepiece, a proper seal has been obtained.

(2) If faceseal air leakage is detected, reposition respirator on face and/or readjust

tension of straps to eliminate air leakage.

160

Cleaning and Storage

Cleaning is recommended after each use. Do not clean with solvents. Cleaning with

solvents may degrade some respirator components and reduce respirator effectiveness.

Inspect all respirator components before each use to ensure proper operating condition.

(1) Remove cartridges and/or filters.

(2) Clean facepiece (excluding filters and cartridges), Respirator Wipes (not to be used

as the only method of cleaning) or by immersing in warm cleaning solution, water

temperature not to exceed 120˚F, and scrub with soft brush until clean. Add neutral

detergent if necessary. Do not use cleaners containing lanolin or other oils.

(3) Disinfect facepiece by soaking in a solution of quaternary ammonia disinfectant or

sodium hypochlorite or other disinfectant.

(4) Rinse in fresh, warm water and air dry in non-contaminated atmosphere.

(5) The cleaned respirator should be stored away from contaminated areas when not in

use.

3.3.2 Eye protection

Splash goggles or safety glasses.

Figure C.9 3M Nuvo Safety Glasses with Clear Anti-Fog Lens

Product information

Manufacturer: 3M Occupational Health and Environmental Safety Division (OH &

ESD), USA.

Product number: 11411-00000-20

Lens Color: Clear - General purpose with impact protection and maximum visibility.

Compliance: ANSI Z87.1-2010 & CSA Z94.3 Certified.

Classic look of a dual-lens. Tough Polycarbonate Lenses Absorb 99.9% of UV. Integral

side shields provide excellent profile protection. Brow bar provides protection from above

and cushion on impact. Soft, Universal-Fit nosepiece for maximum comfort and

161

adjustability. Temples offer adjustability and soft tips for all-day comfort. Anti-Fog Hard

Coat guards against fogging, scratching, static and chemical attack. Lightweight at 1.2 oz.

Using guideline

Wearing the protective eyewear when working with TEG wastewater is a must.

Figure C.10 Worker Wearing Safety Glasses

The glass‘s frames are adjustable to adapt with various users

3.3.3 Hand protection

Nitrile gloves

Figure C.11 Nitrile Gloves

Product information

Manufacturer: Ansell Corporate, Thailand

Material: Nitrile

Product No.: 92-600

Size: Medium

Color: Green

Physical properties: Length 240 mm, average palm width 96 mm, Palm thickness single

wall 0.12 mm.

162

Features: Outstanding Chemical resistance, excellent puncture resistance, easy

donning and strong Grip.

Using guideline

Wearing the nitrile gloves when working with TEG wastewater is a must.

Figure C.12 Wearing Nitrile Glove

Use one time only. This is disposable product

163

Appendix D

Details of Calculations

164

Appendix D.1 Calculations of Membrane Surface Area (outside-in operation)

Bench scale hollow fiber membrane distillation module

Outer diameter of fiber = 0.00203 m

Fiber length = 0.4 m

No. of fibers = 100

Surface Area = 100 x 0.4 x (π x 0.00203)

= 0.255 m2

Pilot scale hollow fiber membrane distillation module

Outer diameter of fiber = 0.00203 m

Fiber length = 1.008 m

No. of fibers = 306

Surface Area = 306 x 1.008 x (π x 0.00203)

= 1.967 m2

Appendix D.2 Calculations of Permeate Flux (outside-in operation)

PWF at 70oC of 0.45 µm bench scale hollow fiber membrane

Initial water level = 88 mm

1-hour water level = 76 mm

1 mm of feed tank ≅ 66.67 g

Water density at 70oC = 977.36 mg/L

Permeate flux = ((88-76) x 66.67)/(0.255 x 977.63)

= 3.14 kg/(m2.h)

Appendix D.3 Thermal Efficiency Calculations

The balance of heat energy is expressed as below

𝑄𝑚 = 𝑄𝑐 + 𝑄𝑣

Where Qm, Qv, Qc are total heat transfer in MD process, heat transfer by mass transfer and

heat transfer by conduction of membrane respectively

𝑄𝑣 = 𝐽𝑤 × ∆𝐻𝑣,𝑤

= 1.94 𝑘𝑔 𝑚2. ℎ⁄ × 2270 𝑘𝐽 𝑘𝑔⁄

= 4403.8 𝑘𝐽 𝑚2. ℎ⁄

𝑄𝑐 =𝑘𝑚

𝛿× (𝑇𝑚𝑓 − 𝑇𝑚𝑝)

=0.14 𝑊 𝑚.𝐾⁄

0.00048 𝑚× (67.7 − 66.6)

165

= 320.8 𝑘𝐽 𝑚2. ℎ⁄

Thermal efficiency, η

η =Qv

Qv+Qc× 100

=4403.8

4403.8 + 320.8× 100

= 93.2 %

𝑄𝑐 is considered as heat loss.

Appendix D.4 Flow Simulation and Heat Transfer Calculation

Pilot Scale Membrane Distillation Membrane Module

Velocity Distribution

Initially, the pilot scale hollow fiber membrane distillation was designed with the length of

1,105 mm and the diameter of flow channel was 47.6 mm. Whereas, the diameter of inlet

pipe was 23 mm. With the inlet flow rate of 60 L/min, the flow velocity in the inlet pipe

was 2.4 L/min, meanwhile the theoretical average value inside the module was 0.16 m/s.

As soon as the high fluid velocity exposed to the hollow fiber at the module’s entrance, it

would damage the fiber. Besides, the hollow fibers also acted as a barrier of fluid entrance.

The flow simulation was done to simulate the characteristic of liquid flow inside the

module (local velocity, local flow direction). The boundary condition was set at: inlet flow

of 1 L/s (60 L/min) and outlet flow at atmospheric pressure. The flow temperature was

70oC with the fluid density of 977.84 kg/m3.

The simulated results by using Solidworks software were shown in Table D.1 and Figure

D.1.

Table D.1 Computational Flow Simulation Results

Name Unit Value Delta

Minimum Velocity m/s -1.030 0.0105

Average Velocity m/s 0.164 5.1687 x 10-6

Maximum Velocity m/s 3.350 0.0036

Maximum Reynolds number 394,700.3 425.9

Average Reynolds number 19,371.7 0.609

Minimum Reynolds number -121,295.7 1235.6

166

The simulated result show that there were reverse flow directions at some positions in the

module as could be seen clearly in Figure D.1. This was due to not well designed module

configuration. However, the average Reynolds number achieved in modeling ( 19,371.7)

was not so different with the theoretical value (18,930.4). This good correlation indicated

that the flow inside the module was under turbulent condition.

Figure D.2 The distribution of velocity inside hollow fiber membrane module

167

Figure D.3 Inlet velocity distribution in hollow fiber module

Reynolds number (Re) (water at 70oC)

Re = (v x Dh x ρ)/ µ

v = 0.16014 m/s

Dh = (4 x Cross section area)/(wetted perimeter)

= (4 x (7.2 x 10-3 – 9.9 x 10-4))/(2 x π x (0.048 + 0.0355))

= 0.04763 m

ρ = 977.36 kg/m3

µ = 0.39 x 10-3 Pa.s

Re = (0.16014 x 0.04763 x 977.36)/( 0.39 x 10-3)

= 18,930.4 (Turbulent flow)

Prandtl number (Pr) (water at 70oC)

Pr = (Cp x µ)/k

Where

cp = specific heat capacity (kJ/(kg.K))

k = thermal conductivity (W/(m.K))

Pr (60oC) = (4066.8 x 0.39 x 10-3)/(0.6627)

= 2.417

168

Nusselt number (Nu)

Nu = 1.86 x (Re x Pr x Dh/L)0.33

Nu (70oC) = 1.86 x (18,930.4 x 2.417 x 0.04763/1.1050)0.33

= 22.7

Heat transfer coefficient of boundary layer (h)

hf (feed) = (Nu x k)/ Dh

= (22.7 x 0.6627)/0.04763

= 316.4 W/(m2.K)

Appendix D.5 Membrane Surface Temperature Calculation

PWF at 70oC of 0.45 µm hollow fiber membrane (permeate flux of 3.14 kg/(m2.h))

Assume Tmf = 60 oC

and Tmp = 55 oC

Using Equation 4.2 Hv = 1425.56 kJ/kg

Substitute Hv to Equation 4.3 and 4.4 yield new values of Tmf and Tmp

Tmf = 66.62 oC

Tmp = 65.49 oC

Repeat the same procedure by using the new Tmf and Tmp



Tmf = 66.59 oC

Tmp = 64.53 oC

Repeat the same procedure by using the new Tmf and Tmp



Tmf = 66.59 oC constant

Tmp = 64.53 oC constant

Finally, membrane surface temperature at: Feed surface 66.59 oC

Permeate surface 64.53 oC

Appendix D.6 Temperature Polarization Coefficient Calculation

PWF at 70oC of 0.45 µm hollow fiber membrane (permeate flux of 3.14 kg/(m2.h))

TPC = (Tmf - Tmp)/(Tbf - Tbp)

= (66.6 – 64.5)/(69-48)

= 0.1

Appendix D.7 Membrane Distillation Coefficient and Resistance Calculation

169

Pure water flux at 70oC (66.6oC feed and 64.5oC permeate surface temperature) of 0.45 µm

Hollow fiber membrane (permeate flux = 3.14 kg/m2.h)

1. Experimental membrane distillation coefficient and resistance (from pure water

flux)

- Vapor pressure at feed side:

vpf = exp (23.1964 −3816.44

T − 46.13)

= exp (23.1964 −3816.44

(66.6 + 273) − 46.13)

= 26,685.3 Pa

- Vapor pressure at permeate side:

Humidity of gas inlet: 49%

Tmp = 64.5oC = 337.5oK → Saturation pressure: 24737.9 Pa

Condensation occurred at the output of fiber → Humidity of gas outlet: 100%

→ vpw = 24737.9 Pa

win = 0.62198 × vpw

101325 − vpw=

0.62198 × 24737.9

101325 − 24737.9= 0.2

Air flow rate: 𝑄𝑎𝑖𝑟 = 25.5 L/min

𝜌𝑎𝑖𝑟 = 1.1126 kg/m3

→ Air mass 𝑚𝑎 = (25.5 × 0.001 × 60) × 1.1126 = 1.70228 kg/h

Total amount of water vapour at the permeate side (membrane surface of 0.255 m2):

𝑊 = win +𝐽 × 0.255

𝑚𝑎= 0.2 +

3.14 × 0.255

1.70228= 0.671271

vpp =W × P

W + 0.622=

0.671271 × (3000 Pa)

0.671271 + 0.622= 1557.147 Pa

- Coefficient:

Bw =3.14 (kg m2h⁄ )

(26,685.3 − 1557.147 ) × 3600 (s h⁄ )= 3.471 × 10−8 (s m⁄ )

- Resistance:

Rw = 1

Bw=

1

3.471 × 10−8= 2.88 × 107 (m s⁄ )

Rw contains membrane resistance and boundary layer resistance.

170

Boundary layer resistance: Rbf and Rbp

Feed side

Bulk feed temperature: Tbf = 68oC → vpbf = 28,384.9 Pa

Feed surface membrane temperature: Tfs = 66.6oC → vpfs = 26,685.3 Pa

Rbf =(vpbf − vpfs)

J=

(28,384.9 − 26,685.3) × 3600 s h⁄

3.14= 1.95 × 106 (m s⁄ )

Permeate side

Humidity: 49%

Tbp = 48oC = 321oK → Saturation pressure: 11339.2 Pa

vpw = 0.49 × 11339.2 = 5556.21 Pa

win = 0.62198 × vpw

101325 − vpw=

0.62198 × 5556.21

101325 − 5556.21= 0.03609


𝜌𝑎𝑖𝑟 = 1.1126 kg/m3

→ Air mass 𝑚𝑎 = (25.5 × 0.001 × 60) × 1.1126 = 1.70228 kg/h


𝑊 = win +𝐽 × 0.255

𝑚𝑎= 0.03609 +

3.14 × 0.255

1.70228= 0.506

vpp =W × P

W + 0.622=

0.506 × (3000 Pa)

0.506 + 0.622= 1346.41 Pa

Rbp =(vpps − vpbp)

J=

(1555.73 − 1346.41 ) × 3600 s h⁄

3.14= 2.4 × 105 (m s⁄ )

Thus

Rbf + Rbp = 1.95 × 106 + 2.4 × 105 = 2.19 × 106 (m s⁄ )

Membrane resistance

R𝑚 = R𝑤 − (Rbf + Rbp) = 2.88 × 107 − 2.19 × 106 = 26.61 × 106 (m s⁄ )

2. Theoretical membrane distillation coefficient and resistance

For Mass Transfer Mechanism: Knudsen and Molecular diffusion combined

- Coefficient:

171

𝐵𝑤 = [3𝜏𝛿

2휀𝑟(

𝜋𝑅𝑇

8𝑀𝑤)

12⁄

+𝜏𝛿𝑃𝑎𝑅𝑇

휀𝑃𝐷𝑀𝑤]

−1

Where

PD = (1.895 × 10−5) × T2.072

= (1.895 × 10−5) × T2.072

= (1.895 × 10−5) × (273 + 66.6)2.072

= 3.325 (Pa m2 s⁄ )

𝐵𝑤 = [3 × 2 × 480 × 10−6

2 × 0.8 × 0.225 × 10−6(

𝜋 × 8.314 × 339.6

8 × 0.018)

12⁄

+2 × 480 × 10−6 × 101325 × 8.314 × 339.6

0.8 × 3.325 × 0.018]

−1

𝐵𝑤 = 1.295 × 10−7(s m⁄ )

- Resistance:

Rw = 1

Bw=

1

1.295 × 10−7= 7.72 × 106 (m s⁄ )

Appendix D.8 Boundary Layer Resistance Calculations

Bulk feed vapor pressure: vpbf = 28,384.9 Pa

Feed surface vapor pressure: vpfs = 26,685.3 Pa

At feed solution with initial TEG concentration of 10%, the permeate flux was 2.61

kg/m2.h.

Feed boundary layer resistance:


J=

(28,384.9 − 26,685.3) × 3600 s h⁄

2.61= 2.34 × 106 (m s⁄ )

Bulk permeate vapor pressure: vppf = 1557.147 Pa

Permeate surface vapor pressure: vpps = 1346.41 Pa

Rbp =(vpps − vpbp)

J=

(1555.73 − 1346.41 ) × 3600 s h⁄

2.61= 2.9 × 105 (m s⁄ )

Total boundary layer resistance:

Rb = Rbf + Rbp = 2.34 × 106 + 2.9 × 105 = 2.63 × 106 (m/s)

Appendix D.9 Fouling analysis

Bench scale hollow fiber membrane SGMD configuration 0.45 𝜇m

Bulk feed temperature: Tbf = 68oC

172

Feed surface membrane temperature: Tfs = 66.6oC

Bulk permeate temperature: Tbp = 48oC

Permeate surface membrane temperature: Tfs = 64.5oC

Fouling flux: J = 2.09 kg/m2h.

Total resistance

J = Bt(vpf − vpp)

Thus,

Bt =J

(vpf − vpp)=

2.09 (kg m2h⁄ )

(26,685.3 − 1557.147 ) × 3600 (s h⁄ )

= 2.31 × 10−8 (s m⁄ )

- Resistance:

R𝑡 = 1

B𝑡=

1

2.31 × 10−8= 4.33 × 107 (m s⁄ )

Boundary layer resistance: Rbf and Rbp

Feed side

Bulk feed temperature: Tbf = 68oC → vpbf = 28,384.9 Pa

Feed surface membrane temperature: Tfs = 66.6oC → vpfs = 26,685.3 Pa


J=

(28,384.9 − 26,685.3) × 3600 s h⁄

2.09= 2.93 × 106 (m s⁄ )

Permeate side

Humidity: 49%

Tbp = 48oC = 321oK → Saturation pressure: 11339.2 Pa

vpw = 0.49 × 11339.2 = 5556.21 Pa

win = 0.62198 × vpw

101325 − vpw=

0.62198 × 5556.21

101325 − 5556.21= 0.03609


𝜌𝑎𝑖𝑟 = 1.1126 kg/m3

→ Air mass 𝑚𝑎 = (25.5 × 0.001 × 60) × 1.1126 = 1.70228 kg/h


𝑊 = win +𝐽 × 0.255

𝑚𝑎= 0.03609 +

3.14 × 0.255

1.70228= 0.506

vpp =W × P

W + 0.622=

0.506 × (3000 Pa)

0.506 + 0.622= 1346.41 Pa

Rbp =(vpbp − vpps)

J=

(1557.14 − 1346.41 ) × 3600 s h⁄

2.09= 3.63 × 105 (m s⁄ )

173

Thus

Rbf + Rbp = 2.93 × 106 + 3.63 × 105 = 3.29 × 106 (m s⁄ )

Fouling resistance:

R𝑓 = R𝑡 − (R𝑚 + Rbf + Rbp) = 4.33 × 107 − (26.61 × 106 + 3.29 × 106)

= 13.37 × 106 (m s⁄ )

Recoverable fouling, Rr

The pure water flux after rising with DI water was 2.88 kg/m2.h

R𝑡2 = 3.14 × 107 m/s

Rbf2 = 2.125 × 106 m/s Rbp2 = 2.63 × 105 (m s⁄ )

R𝑓2 = R𝑡2 − (R𝑚 + Rbf2 + Rbp2) = 2.4 × 106 (m s⁄ )

Thus,

R𝑟 = 13.37 × 106 − 2.4 × 106 = 10.97 × 106(m s⁄ )

Reversible fouling, Rre

The pure water flux after rising with chemical agents and DI water was 3.05 kg/m2.h

R𝑡3 = 2.97 × 107 m/s

Rbf3 = 2.006 × 106 m/s Rbp3 = 2.49 × 105 (m s⁄ )

R𝑓3 = R𝑡2 − R𝑡3 − (Rbf3 + Rbp3) = 0.79 × 106 (m s⁄ )

Therefore

R𝑟𝑒 = R𝑓2 − R𝑓3 = 2.4 × 106 − 0.79 × 106 = 1.62 × 106 (m/s)

Irreversible fouling, Rirre

From

R𝑓 = R𝑟 + R𝑟𝑒 + R𝑖𝑟𝑟𝑒

Thus,

R𝑖𝑟𝑟𝑒 = R𝑓 − (R𝑟 + R𝑟𝑒)

= 13.37 × 106 − 10.97 × 106 − 1.62 × 106

= 7.85 × 105 (m/s)

174

Appendix E

Experimental Results

175

1. Bench scale hollow fiber membrane distillation (0.45 µm)

Table E.1 Pure Water Flux at 70oC feed, 2.4 L/min feed flow rate, 3.1 m/s sweeping

gas

Time (h) Feed level (cm) Permeate flux (kg/m2.h) Energy ratio (kW/kg)

0 9.6 - -

1 8.9 1.83 1.27

2 8.2 1.83 1.33

3 7.5 1.83 1.44

Average 1.83 1.35

Table E.2 Pure Water Flux (at 70oC feed, 2.4 L/min feed flow rate, 3.6 m/s sweeping

gas)


0 11.2 - -

1 10.4 2.09 1.27

2 9.6 2.09 1.21

3 8.8 2.09 1.46

Average 2.09 1.31


gas)


0 12.6 - -

1 11.6 2.61 1.11

2 10.6 2.61 1.04

3 9.6 2.61 1.07

Average 2.61 1.07


gas)


0 8.8 - -

1 7.6 3.14 1.06

2 6.4 3.14 1.06

3 5.2 3.14 1.16

Average 3.14 1.09

176


gas)


0 15.7 - -

1 14.7 2.61 1.59

2 13.6 2.88 1.06

3 12.6 2.61 1.24

Average 2.70 1.30


gas)


0 17.9 - -

1 17.4 1.31 1.39

2 16.9 1.31 1.39

3 16.4 1.31 1.42

Average 1.31 1.40


gas)


0 16.4 - -

1 15.8 1.57 1.37

2 15.2 1.57 1.17

3 14.6 1.57 1.29

Average 1.57 1.27


gas)


0 14.6 - -

1 13.9 1.83 1.46

2 13.2 1.83 1.24

3 12.5 1.83 1.31

Average 1.83 1.34


gas)


0 13.2 - -

1 12.5 2.09 1.35

2 11.8 1.83 1.42

3 0 1.83 0.63

Average 1.92 1.13

177


gas)


0 10.2 - -

1 9.6 1.57 1.85

2 0 1.57 0.82

3 0 1.57 0.82

Average 1.57 1.17


gas)


0 9.4 - -

1 9.2 0.52 2.86

2 9 0.52 2.56

3 8.8 0.52 2.64

Average 0.52 2.69


gas)


0 11.7 - -

1 11.4 0.78 2.18

2 11.1 0.78 1.88

3 10.8 0.78 1.88

Average 0.78 1.98


gas)


0 10.8 - -

1 10.4 1.05 1.54

2 10 1.05 1.65

3 9.6 1.05 1.58

Average 1.05 1.59

178


gas)


0 9.6 - -

1 9.2 1.05 1.78

2 8.8 1.05 1.89

3 8.4 1.05 1.11

Average 1.05 1.59


gas)


0 8.4 - -

1 8 1.05 1.99

2 7.6 1.05 1.24

3 7.2 1.05 1.24

Average 1.05 1.49

Table E.16 Pure Water Flux (at 70oC feed, 2 L/min feed flow rate, 3.1 m/s sweeping

gas)


0 9.6 - -

1 8.9 1.83 1.38

2 8.2 1.83 1.50

3 7.5 1.83 1.42

Average 1.83 1.43


gas)


0 14.9 - -

1 14.1 2.09 1.29

2 13.25 2.22 1.21

3 12.5 1.96 1.43

Average 2.09 1.31

179


gas)


0 12 - -

1 11 2.61 1.13

2 10 2.61 1.23

3 9 2.61 1.13

Average 2.61 1.16


gas)


0 11.7 - -

1 10.65 2.75 1.22

2 9.5 3.01 1.02

3 8.45 2.75 1.31

Average 2.83 1.18


gas)


0 16.9 - -

1 16 2.35 1.48

2 15.1 2.35 1.37

3 14.2 2.35 1.58

Average 2.35 1.42


gas)


0 11.1 - -

1 10.6 1.31 1.27

2 10.1 1.31 1.39

3 9.6 1.31 1.33

Average 1.31 1.33

180


gas)


0 9.6 - -

1 9 1.57 1.37

2 8.4 1.57 1.27

3 7.8 1.57 1.32

Average 1.57 1.32


gas)


0 16.3 - -

1 15.6 1.83 1.31

2 14.9 1.83 1.20

3 14.2 1.83 1.41

Average 1.83 1.31


gas)


0 14.2 - -

1 13.4 2.09 1.23

2 12.6 2.09 1.25

3 11.8 2.09 1.27

Average 2.09 1.25


gas)


0 11.8 - -

1 11.2 1.57 1.65

2 10.6 1.57 1.92

3 9.9 1.83 1.50

Average 1.66 1.69

181


gas)


0 15.1 - -

1 14.9 0.52 2.56

2 14.7 0.52 2.56

3 14.5 0.52 2.56

Average 0.52 2.56


gas)


0 14.3 - -

1 14 0.78 1.93

2 13.7 0.78 1.88

3 13.4 0.78 1.88

Average 0.78 1.90


gas)


0 13.3 - -

1 12.9 1.05 1.58

2 12.5 1.05 1.61

3 12.1 1.05 1.54

Average 1.05 1.58


gas)


0 11.5 - -

1 11.1 1.05 1.86

2 10.7 1.05 1.74

3 10.3 1.05 1.74

Average 1.05 1.78

182


gas)


0 10.5 - -

1 10.1 1.05 2.02

2 9.7 1.05 1.95

3 9.3 1.05 2.14

Average 1.05 2.04


gas)


0 15.6 - -

1 14.9 1.83 1.40

2 14.2 1.83 1.33

3 13.7 1.31 1.84

Average 1.66 1.52


gas)


0 13.5 - -

1 12.7 2.09 1.23

2 11.9 2.09 1.21

3 11.1 2.09 1.29

Average 2.09 1.24


gas)


0 10.4 - -

1 9.5 2.35 1.18

2 8.5 2.61 1.14

3 7.6 2.35 1.32

Average 2.44 1.21

183


gas)


0 16.1 - -

1 15 2.88 1.19

2 14 2.61 1.07

3 13 2.61 1.10

Average 2.70 1.12


gas)


0 13 - -

1 12 2.61 1.32

2 11 2.61 1.18

3 10 2.61 1.21

Average 2.61 1.24

Table E.36 Synthetic TEG Wastewater Investigation of Hollow fiber Membrane

(10%)

Time

(h)

Permeate flux

(kg/m2.h)

TEG

Concentration (%)

Energy ratio

(kW/kg)

1 2.61 11.3 1.30

2 2.61 12.2 1.12

3 2.61 13.2 1.33

4 2.61 14.4 1.22

5 2.61 15.9 1.24

6 2.61 17.7 1.28

7 2.61 19.9 1.12

8 2.61 22.7 1.31

9 2.09 25.6 1.53

10 1.83 28.9 1.75

11 2.09 33.8 1.40

12 2.09 44.6 1.02

184


(30%)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 1.57 32.6 2.11

2 2.09 33.1 1.38

3 1.83 34.9 1.79

4 2.09 37.2 1.42

5 2.35 40.2 1.36

6 2.09 43.3 1.32

7 1.83 46.4 1.64

8 2.09 50.6 1.38

9 2.09 55.6 1.42

10 1.31 59.2 2.17

12 1.57 60.4 1.86


(60%)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 0.52 60.8 4.76

2 1.05 62.5 2.72

3 1.05 64.3 2.38

4 1.05 66.2 2.46

5 1.05 68.2 2.42

6 1.05 70.3 2.38

7 0.78 72.0 2.92

8 0.78 73.8 3.12

9 0.52 75.0 4.69

11 0.52 77.6 3.21

12 0.52 78.9 4.46

13 0.52 80.4 4.24

23 0.00 90.5 4.41

Table E.39 Continuously-fed Synthetic TEG Wastewater Investigation of Hollow

fiber Membrane

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 2.61 10.7 1.13

2 2.61 11.2 1.06

3 2.61 12.1 1.43

4 2.61 13.2 1.30

5 2.61 14.5 1.27

6 2.61 16.2 1.28

7 2.61 18.2 1.21

185

8 2.61 20.8 1.31

8 15.0

9 2.61 1.19

16 2.09 29.7 1.62

16 19.7

17 2.61 21.1 1.06

18 2.61 22.7 1.06

19 2.61 24.6 1.33

20 2.35 24.6 1.19

21 2.09 26.4 1.47

22 2.35 28.7 1.19

23 1.83 30.8 1.55

24 2.09 33.6 1.38

24 22.4

25 2.35 23.8 1.27

26 2.35 25.5 1.22

27 2.35 27.3 1.26

28 2.35 29.5 1.17

29 2.35 29.5 1.46

30 2.35 32.0 1.19

31 2.09 34.7 1.53

32 2.09 37.8 1.32

32 24.8

33 2.35 1.34

40 2.09 53.2 1.45

Table E.40 Continuously-fed Synthetic TEG Wastewater Investigation of Hollow

fiber Membrane (209 hours)

Time

(h)

Permeate flux

(kg/m2.h)

Energy consumption

(KW/h)

Energy ratio

(kW/kg)

TEG

Concentration

(%)

1 9.69

2 2.35 0.74 1.2

3 2.88 0.83 1.1

4 2.88 0.80 1.1

5 2.61 0.84 1.3

6 2.09 0.77 1.4

7 2.61 0.82 1.2

8 2.61 0.79 1.2

9 2.35 0.80 1.3

10 2.35 0.87 1.4

11 2.61 0.99 1.5

13 2.61 0.82 1.2 17.25

14 2.61 0.87 1.3

186

15 2.61 0.78 1.2

16 2.61 0.85 1.3

17 2.61 0.87 1.3

18 2.61 0.81 1.2

19 2.61 0.83 1.3

20 2.61 0.82 1.2

22 2.61 0.82 1.2

23 2.61 0.84 1.3

25 2.35 0.80 1.3 31.24

26 2.61 0.76 1.1

27 2.61 0.76 1.1

28 2.61 0.90 1.3

29 2.09 0.70 1.3

30 2.09 0.62 1.2

31 3.14 0.81 1.0

32 2.09 0.74 1.4

33 2.09 0.74 1.4

35 2.09 0.78 1.5

36 2.61 0.75 1.1

37 2.61 0.76 1.1

38 2.35 0.78 1.3

39 2.35 0.75 1.2

43 2.29 0.76 1.3

44 1.83 0.78 1.7

45 1.83 0.79 1.7

47 2.09 0.80 1.5

48 2.09 0.80 1.5

49 2.61 0.78 1.2

50 2.35 0.81 1.4

51 2.35 0.72 1.2

56 2.09 0.75 1.4

57 1.83 0.83 1.8

58 1.83 0.68 1.5

59 1.57 0.80 2.0

61 1.31 0.72 2.1 41.87

62 1.83 0.78 1.7

63 2.35 0.80 1.3

64 2.35 0.81 1.3

65 2.35 0.70 1.2

66 1.57 0.80 2.0

67 1.57 0.73 1.8

69 1.83 0.77 1.6

70 2.09 0.73 1.4

187

71 2.09 0.65 1.2

80 1.96 0.79 1.6 56.24

81 1.57 0.60 1.5

82 1.31 0.66 2.0

83 1.31 0.60 1.8

84 1.05 0.62 2.3

85 1.57 0.66 1.7

87 1.83 0.63 1.3

88 1.83 0.83 1.8

90 1.83 0.73 1.6

92 1.83 0.66 1.4

93 1.57 0.79 2.0

95 2.09 0.80 1.5 71.18

106 1.38 0.66 1.9

107 1.31 0.65 1.9

108 1.31 0.63 1.9

109 1.05 0.58 2.2

111 0.65 0.59 3.6

112 0.78 0.58 2.9

113 0.26 0.53 7.9

115 2.09 0.79 1.5

116 1.83 0.69 1.5

118 1.70 0.73 1.7

131 1.15 0.64 2.2

132 0.26 0.64 9.7

133 0.52 0.60 4.5

135 2.09 0.67 1.3 77.38

136 1.05 0.64 2.4

137 2.09 0.68 1.3

155 1.00 0.62 2.4

156 0.13 0.69 20.5

158 2.09 0.66 1.2 98.01

160 1.57 0.67 1.7

161 1.31 0.66 2.0

162 2.35 0.93 1.6

163 1.57 0.69 1.7

165 1.05 0.66 2.5

176 0.76 0.63 3.3

177 0.26 0.41 6.2

178 0.26 0.51 7.6

179 0.26 0.47 7.0

180 0.26 0.49 7.4

182 1.57 0.63 1.6

183 1.57 0.66 1.7

188

186 1.05 0.63 2.4

187 1.05 0.70 2.6

188 1.05 0.60 2.2

189 1.05 0.61 2.3

200 0.55 0.56 4.0

201 0.52 0.52 3.9

202 0.52 0.53 3.9

203 0.26 0.48 7.3

205 1.31 0.56 1.7

206 1.05 0.72 2.7

207 1.05 0.65 2.5

208 1.05 0.76 2.8

209 1.05 0.69 2.6

Table E.41 Real Wastewater Experiment of hollow fiber membrane (1st batch)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 2.61 9.69 1.19

2 2.61 10.38 1.22

3 2.35 11.10 1.27

4 2.35 11.91 1.17

5 2.35 12.86 1.47

6 2.35 13.98 1.47

7 2.09 15.14 1.39

8 2.35 16.71 1.31

Table E.42 Real Wastewater Experiment of hollow fiber membrane (2nd batch)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 2.61 9.8 1.21

2 2.61 10.38 1.25

3 2.09 11.10 1.45

4 2.09 11.91 1.25

5 2.09 12.86 1.58

6 2.35 13.98 1.27

7 2.35 15.14 1.07

8 2.35 16.71 1.26

189

Table E.43 Real Wastewater Experiment of hollow fiber membrane (3rd batch)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 2.61 9.70 1.29

2 2.35 10.31 1.30

3 2.35 11.01 1.40

4 2.35 11.82 1.20

5 2.35 12.75 1.20

6 2.35 13.84 1.30

7 2.35 15.14 1.20

8 2.35 16.71 1.30

Table E.44 Real Wastewater Experiment of hollow fiber membrane (4th batch)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 2.35 9.69 1.25

2 2.35 10.31 1.30

3 2.35 11.01 1.30

4 2.35 11.82 1.30

5 2.35 12.75 1.20

6 2.35 13.84 1.30

7 2.35 15.14 1.20

8 2.09 16.52 1.30

Table E.45 Real Wastewater Experiment of hollow fiber membrane (5th batch)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

1 2.35 9.69 1.29

2 2.35 10.31 1.3

3 2.35 11.01 1.3

4 2.35 11.82 1.3

5 2.35 12.75 1.2

6 2.35 13.84 1.3

7 2.09 14.98 1.4

8 2.09 16.33 2.3

Table E.46 Real Wastewater Experiment of hollow fiber membrane (40 hours)

Time

(h)

Permeate flux

(kg/m2.h)

TEG Concentration

(%)

Energy ratio

(kW/kg)

0 - 9.69 -

1 2.35 1.2

7 - 16.90 -

8 2.09 25.13 1.3

8 - 20.23 -

9 2.35 21.52 1.1

190

10 2.35 22.99 1.3

11 2.09 24.47 1.5

12 2.09 26.16 1.4

13 2.09 28.10 1.4

14 2.09 30.50 1.3

15 2.09 33.20 1.3

16 2.09 35.70 1.3

16 - 27.56 -

17 2.35 29.32 1.0

18 2.35 31.32 1.2

19 2.09 33.34 1.6

20 2.09 35.64 1.3

21 2.09 38.28 1.2

22 1.83 40.93 1.4

23 1.83 43.98 1.5

24 1.83 47.52 1.5

24 - 30.44 -

25 2.09 0.6

32 1.57 51.62 1.8

32 - 34.81 -

33 2.09 36.77 1.2

34 2.09 38.97 1.1

35 2.09 41.44 1.3

36 1.83 43.88 1.6

37 1.83 46.90 1.6

38 1.83 49.20 1.6

39 1.57 52.10 1.7

40 1.57 55.55 1.7

Table E.47 Fouling Investigation after Cleaning with DI Water (Batch Experiment)


0 10.6 - -

1 9.5 2.88 1.10

2 8.4 2.88 1.17

3 7.3 2.88 1.15

Average 2.88 1.14

191

Table E.48 Fouling Investigation after Cleaning with Chemical Agents (Batch

Experiment)


0 9 - -

1 7.8 3.14 1.06

2 6.6 3.14 0.94

3 5.5 2.88 1.23

Average 3.05 1.08

Table E.49 Fouling Investigation after Cleaning with DI Water (Continuously-fed

Experiment)


0 13.7 - -

1 12.6 2.88

1.15

2 11.5 2.88

3 10.4 2.88

Average 2.88 1.15

Table E.49 Fouling Investigation after Cleaning with Chemical Agents (Continuously-

fed Experiment)


0 15.3 - -

1 14.1 3.14 1.01

2 12.9 3.14 1.06

3 11.8 2.88 1.18

Average 3.05 1.08

2. Pilot Scale Study

Table E.51 Pure Water Flux Investigation (feed flow rate 60 L/min, feed temperature

70oC)

Sweeping

Gas Inlet

Velocity

(m/s)

First Experiment

Duplicated

Experiment Average

Permeate

flux

(kg/m2.h)

Energy

ratio

(kW/kg)

Permeate

flux

(kg/m2.h)

Energy

ratio

(kW/kg)

Permeate

flux

(kg/m2.h)

Energy

ratio

(kW/kg)

4.29 1 0.586 0.95 0.577 0.98 0.58

4.98 1.19 0.552 1.19 0.541 1.19 0.55

5.25 1.55 0.456 1.51 0.460 1.53 0.46

6.06 1.59 0.486 1.66 0.465 1.63 0.48

6.76 1.75 0.495 1.83 0.459 1.79 0.48

7.57 1.75 0.548 1.89 0.493 1.82 0.52

8.07 1.91 0.518 1.97 0.502 1.94 0.51

192

8.58 1.89 0.553 1.96 0.538 1.93 0.55

9.09 1.91 0.578 1.86 0.571 1.89 0.57

9.59 1.95 0.590 - - 1.95 0.59

10.09 1.99 0.604 - - 1.99 0.60

10.60 1.99 0.627 - - 1.99 0.63

193

Table E.52 Synthetic TEG Experiment (10%)

Time

(h)

First Experiment Duplicated Experiment Average

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

TEG

concentration

(%)

1 1.91 0.52 1.83 0.53 1.87 0.53 10.8

2 1.75 0.56 1.87 0.53 1.81 0.54 11.7

3 1.75 0.55 1.75 0.57 1.75 0.56 12.7

4 1.67 0.59 1.71 0.58 1.69 0.58 13.8

5 1.63 0.60 1.83 0.53 1.73 0.57 15.2

6 1.59 0.59 1.67 0.58 1.63 0.59 16.8

7 1.67 0.59 1.59 0.60 1.63 0.60 18.8

8 1.67 0.60 1.87 0.51 1.77 0.56 21.4


Time

(h)


Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

TEG concentration

(%)

1 1.87 0.51 1.75 0.54 1.87 0.53 21.5

2 1.67 0.58 1.67 0.60 1.81 0.54 23.1

3 1.67 0.59 1.67 0.58 1.75 0.56 25.0

4 1.67 0.59 1.67 0.59 1.69 0.58 27.2

5 1.63 0.61 1.63 0.59 1.73 0.57 29.7

6 1.59 0.61 1.51 0.63 1.63 0.59 32.6

7 1.59 0.61 1.79 0.57 1.63 0.60 36.5

8 1.51 0.65 1.47 0.63 1.77 0.56 40.8

194


Time

(h)


Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

TEG concentration

(%)

1 1.67 0.57 1.71 0.57 1.69 0.57 32.1

2 1.67 0.59 1.63 0.60 1.65 0.59 34.5

3 1.63 0.59 1.59 0.61 1.61 0.60 37.1

4 1.51 0.64 1.47 0.64 1.49 0.64 40.0

5 1.55 0.62 1.71 0.59 1.63 0.60 43.5

6 1.51 0.64 1.59 0.60 1.55 0.62 47.6

7 1.39 0.68 1.55 0.62 1.47 0.65 52.0

8 1.39 0.71 1.47 0.66 1.43 0.69 57.4


Time

(h)


Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

TEG concentration

(%)

1 1.51 0.63 1.63 0.60 1.57 0.61 42.59

2 1.47 0.65 1.55 0.63 1.51 0.64 45.42

3 1.43 0.68 1.43 0.68 1.43 0.68 48.46

4 1.43 0.67 1.43 0.66 1.43 0.66 51.95

5 1.35 0.71 1.55 0.62 1.45 0.67 56.04

6 1.35 0.71 1.39 0.68 1.37 0.69 60.55

7 1.23 0.77 1.35 0.69 1.29 0.73 65.51

8 1.23 0.75 1.35 0.72 1.29 0.73 71.37

195


Time

(h)


Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

Permeate flux

(kg/m2.h)

Energy ratio

(kW/kg)

TEG

concentration

(%)

1 1.10 0.86 1.27 0.73 1.19 0.79 62.89

2 1.15 0.83 1.27 0.77 1.21 0.80 66.14

3 1.15 0.82 1.19 0.79 1.17 0.80 69.63

4 0.99 0.94 1.19 0.81 1.09 0.87 73.23

5 0.99 0.94 1.19 0.81 1.09 0.87 77.23

6 0.99 0.94 1.19 0.80 1.09 0.87 81.70

7 0.76 1.22 0.72 1.25 0.74 1.23 84.99

8 0.38 2.48 0.52 1.78 0.45 2.13 88.04

196

Table E.57 Synthetic TEG (10%) Experiment at low gas flow rate (34.5 L/min)

Time

(h)

Feed Level

(cm)

Permeate flux

(kg/m2.h)

TEG concentration

(%)

1 46.5 0.83 10.6

2 44.4 0.83 11.0

3 42.3 0.83 11.5

4 40.1 0.87 12.0

5 36.2 0.78 13.1

6 34.3 0.76 13.7

7 32.3 0.80 14.4

8 30.4 0.76 15.1

9 28.6 0.72 15.8

10 26.6 0.80 19.2


Time

(h)

Feed Level

(cm)

Permeate flux

(kg/m2.h)

TEG concentration

(%)

1 42 0.76 28.0

2 39.8 0.87 29.3

3 37.9 0.76 30.5

4 35.6 0.91 32.2

5 33.5 0.83 33.8

6 31.5 0.80 35.5

7 29.7 0.72 37.2

8 27.7 0.80 39.3

9 25.8 0.76 41.6

13 18.2 0.72 44.6


Time

(h)

Feed Level

(cm)

Permeate flux

(kg/m2.h)

TEG

concentration

(%)

Energy ratio

(kW/kg)

1 42.6 0.83 44.8 0.40

2 41.0 0.64 46.2 0.51

3 39.2 0.72 48.0 0.43

4 37.4 0.72 49.9 0.44

5 35.5 0.76 52.1 0.46

6 33.7 0.72 54.4 0.48

7 30.2 0.70 59.4 0.57

8 28.7 0.60 61.9 0.59

9 27.0 0.68 64.9 0.50

197

10 25.3 0.68 68.3 0.50

11 23.8 0.60 71.5 0.55


Time

(h)

Feed Level

(cm)

Permeate flux

(kg/m2.h)

TEG

concentration

(%)

Energy ratio

(kW/kg)

1 43.0 0.64 66.6 0.56

2 41.7 0.52 68.4 0.53

3 40.2 0.60 70.5 0.60

4 38.8 0.56 72.7 0.66

5 37.3 0.60 75.1 0.47

6 35.9 0.56 77.5 0.57

7 34.4 0.60 80.3 0.54

8 32.1 0.46 84.9 0.84

9 30.9 0.48 87.5 0.67

10 29.8 0.44 90.1 0.74

11 28.7 0.44 92.8 0.75


Time

(h)

Feed Level

(cm)

Permeate flux

(kg/m2.h)

TEG

concentration

(%)

Energy ratio

(kW/kg)

1 45.1 0.52 77.6 0.60

2 44.0 0.44 79.2 0.71

3 43.0 0.40 80.7 0.80

4 42.0 0.40 82.3 0.73

5 41.1 0.36 83.7 0.82

6 40.4 0.28 84.9 1.08

7 39.7 0.28 86.1 1.00

8 38.5 0.24 88.2 1.55

9 38.0 0.20 89.2 1.51

10 37.5 0.20 90.1 1.51

11 37.0 0.20 91.1 1.48

12 36.0 0.20 93.0 1.58

13 35.6 0.16 93.9 1.98

14 35.5 0.04 98.4 7.02

Table E.62 Real TEG Wastewater Experiment (Batch 1)

Time

(h)

Feed Level

(cm)

Permeate flux

(kg/m2.h)

TEG

concentration

(%)

Energy ratio

(kW/kg)

1 46.0 1.91 10.46 0.52

2 41.2 1.91 11.37 0.51

3 36.5 1.87 12.42 0.53

198

4 31.7 1.91 13.72 0.52

5 27.3 1.75 15.18 0.56

6 23.0 1.71 16.93 0.58

7 18.5 1.79 19.26 0.57

8 14.0 1.79 22.34 0.55

Table E.63 Real TEG Wastewater Experiment (Batch 5)

Time

(h)

Feed Level

(cm)

Permeate

flux

(kg/m2.h)

TEG

concentration

(%)

Energy ratio

(kW/kg)

0 52.1 - 9.69

1 47.6 1.79 10.4 0.55

7 20.6 - - -

8 16.4 0.83 21.5 1.05


Time

(h)

Permeate flux

(kg/m2.h)

TEG concentration

(%)

Energy ratio

(kW/kg)

0 - 9.7 -

1 1.91 10.5 0.50

8 1.55 24.9 0.61

8 - 19.2 -

9 1.55 20.4 0.61

10 1.51 21.7 0.63

11 1.71 23.5 0.58

12 1.71 25.6 0.56

13 1.55 27.8 0.62

14 1.59 30.5 0.60

15 1.51 33.7 0.65

16 1.51 37.5 0.65

16 - 30.9 -

17 1.59 32.9 0.61

18 1.55 35.1 0.63

19 1.55 37.7 0.62

20 1.51 39.8 0.62

21 1.47 42.2 0.65

22 1.39 44.3 0.69

23 1.35 45.1 0.71

24 1.35 47.0 0.71

24 - 39.5 -

25 1.63 42.2 0.59

26 1.47 44.9 0.65

27 1.47 48.0 0.64

28 1.43 51.5 0.67

199

29 1.43 53.4 0.68

30 1.31 55.8 0.73

31 1.19 57.9 0.81

32 1.19 60.1 0.81

32 - 40.7 -

33 1.31 42.8 0.72

40 0.99 69.4 0.95


Time

(h)

TEG concentration

(%)

Permeate flux

(kg/m2.h)

0 19.6 -

1 - 1.91

8 30.5 1.43

8 27.9 -

9 - 1.87

16 50.0 1.35

16 29.9 -

17 - 1.87

24 66.7 1.39

24 48.9 -

25 - 1.83

32 76.2 0.87

32 57.2 -

33 - 1.75

40 82.7 0.83

40 57.0 -

41 - 1.23

48 86.9 0.68

48 64.3 -

49 - 0.91

56 95.2 0.56

56 72.93 -

57 - 0.95

64 89.57 0.52

64 78.8 -

65 - 0.76

72 99.08 0.39

200

Appendix F

Energy Consumption Analysis

201

1. Bench Scale SGMD Introduction

The schematic diagram of bench scale hollow fiber SGMD system is shown in

Figure F.1.

Figure F.1 Bench Scale SGMD System

The equipment that consumed energy of bench scale SGMD system are presented in

Figure F.2

(a) (b) (c) (d)

Figure F.2 Equipment Consume Energy

a. Feed pump b. Heater c. Control box d. Compressor

The energy consumption of electrical heater, feed pump and control box were measured

by a power rotameter

Heater (feed tank)

Model: Seagull - Digital Pro Electric Urn

Company: Thai Stainless Steel Co.,Ltd

Max capacity: 14 L

Permeate tank

Heater

Feed

tank Gas Compressor

MD Module

202

Feed pump

Company: Iwaki Co.

Model: MD-10L-220

Full load power: 35 W/h

Max head: 15 m

Max capacity: 11L/min

The energy consumption of gas compressor was measured theoretically based on full

load power of the machine.

Air compressor

Company: Hitachi

Model: Bebicon – Oil free- 4C35823

Full load power: 3.7 kWh

Motor efficiency: 80%

Free air delivery: 400L/min

The equation to calculate energy consumption of gas compressor is as below:

ECcompressor = 3.7 kWh × (1/ε) × sweeping gas flow rate ×1

400 L/min

Where ε is compressor coefficient (ε = 0.8),

Example: When investigating the PWF at 25.5 L/min

→ Energy consumption for compressor:

EC = 3.7 kWh x (1/0.8) x 25.5 L/min x 1/(400 L/min) = 0.29 kWh

2. Energy Measurement

Energy Consumption

The energy consumption (EC) of bench scale system was calculated as following

EC (kW/h) = ECcompressor + (ECpump + ECHeater )

= ECcompressor + ECobserving from power meter

Where EC (kW/h) is energy consumption of equipment per hour. The EC show the amount

of energy required to maintain the driving force of MD process

Energy Ratio

The energy ratio represented the relationship between energy consumption of whole

system and permeate flow rate of membrane module.

203

Energy ratio (kW kg)⁄ = Energy consumption (kW/h)

Permeate flux (kg m2⁄ .h) × Membrane area (m2)

optimizing membrane distillation process for triethylene...

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