membrane distillation application in purification and process
TRANSCRIPT
Membrane Distillation Application in Purification and Process
Intensification
by
Dinh Le My
A thesis submitted in partial fulfillment of the requirements for the
degree of Master of Science in
Environmental Engineering and Management
Examination Committee: Prof. Chettiyappan Visvanathan (Chairperson)
Prof. Ajit Padmakar Annachhatre
Dr. Loc Nguyen Thai
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
Acknowledgements
I would like to deeply express my appreciation to Prof. Chettiyappan Visvanathan because
of his patient guidance, efficacious suggestion and encouragement throughout the period of
my thesis.
I am grateful to my committee members - Prof. Ajit P. Annachhatre and Dr. Loc Nguyen
Thai, who gave me the frank comments and suggestion to improve my thesis.
I also acknowledge the help from Mr. Paul Jacob for his advice and Mr. Thusitha for
providing technical comments. In addition, the solidarity as well as the sharing of
experiences from members of Prof. Visu's team, especially Mr. Pham Minh Duyen, have
created the motivation to help me complete the thesis in the best way.
I am thankful to EEM staff and technicians. My experiment is smoother and more safety
with the support from them.
I appreciate the enthusiastic help from Food Engineering Laboratory for their chemical and
analytical procedures in food field. Thanks also would like to be sent for Mr. Kaji who
required for supplying glucose liquid from Ajinomoto Company and Sumitomo Company
for supporting the Membrane to my research work.
A special thank is to Royal Thai Government for granting me the Loong Nam Khong Pijai
Scholarship, creating a chance for me to get a master degree at AIT – Thailand.
Last but not least, I would like to say thank you for the mental support and encouragement
from my family.
iii
Abstract
The concept of process intensification (PI) is implemented in food industry, typically on
concentrating glucose liquid. In wastewater treatment, this concept is used for removing
TDS from phenol Industry (sodium sulfate salt). Membrane Distillation process with
hydrophobic membrane was applied as the promising technology. The feed temperature
was chosen at 70 oC for TDS test and 50 oC for glucose test. In TDS test, there is no
significantly difference of flux between DCMD and SGMD configuration when salt
concentration increased from 40 to 450 g/L. The energy ratio consumed in SGMD was
much lower than in DCMD. Fouling resistance did not play an important role in TDS test.
The highest resistance accounted for 45% that was localized in membrane resistance. Zero
percent of irreversible was found after cleaning the membrane with acidic solution.
However, the energy consumption ratio of DCMD system was markedly higher compared
with used energy ratio in SGMD system. Therefore, SGMD is selected as the favorable
configuration in TDS removal test.
The application of MD on concentrating glucose was fruitful in both DCMD and SGMD
configuration. In DCMD, an 83.9% flux reduction was observed after 269 hours operation
due to fouling resistance accounted for 79.5% of total resistance. HF SGMD system
consumed 60 hours to concentrate real glucose liquid with slowly flux reduction. Fouling
resistance in SGMD did not play as a major role in resistance. The 7 times higher flux, 4
times lower time consumption, 9 times lower energy consumption ratio compared to
DCMD make SGMD become an encouraging configuration in process intensification.
iv
Graphical Abstract
v
Table of Contents
Chapter Title Page
Title page i Acknowledgements ii
Abstract iii
Table of Contents v
List of Tables vii
List of Figures ix
List of Abbreviations xi
1 Introduction 1
1.1 Background 1
1.2 Objectives of the Study 2
1.3 Scope of the Study 2
2 Literature Review 4
2.1 Total Dissolve Solid 4
2.1.1 Definition and its source 4
2.1.2 The related environmental problem 5
2.1.3 TDS in industrial wastewater, its measurement 6
2.1.4 Current TDS removal method 7
2.2 Glucose 15
2.2.1 Definition 15
2.2.2 Physical and chemical properties 16
2.2.3 Application of glucose 17
2.2.4 Glucose syrup 18
2.2.5 The current techniques for concentrating glucose 21
2.3 Membrane Technology 22
2.4 Membrane Distillation 23
2.4.1 Membrane distillation application 23
2.4.2 Membrane distillation configuration 24
2.4.3 Membrane characteristics 26
2.4.4 Mechanism of MD transport 28
2.4.5 Operating parameter 36
2.4.6 Fouling and solution 37
2.5 Research Gap 39
41
3 Methodology 41
3.1 Methodology Framework 41
3.1.1 System calibration 41
3.1.2 The research road map 42
3.2 Experimental Set up 42
3.2.1 Hollow fiber membrane 45
3.3.2 Membrane configuration 46
3.2.3 Membrane module 47
3.3 Experimental Procedure 48
3.3.1 Flow rate calibration 48
3.3.2 Temperature calibration 48
3.3.3 Gas flowrate calibration (in SGMD) 48
vi
3.3.4 Temperature polarization 48
3.3.5 System verification 49
3.3.6 The concentration of feeding solution 50
3.3.7 Energy consumption 51
3.4 Parameter Analysis 51
3.4.1 Glucose analysis 51
3.4.2 Sodium sulfate analysis 52
3.5 Membrane Cleaning 52
55
4 Results and Discussion 55
4.1 Membrane Distillation System Calibration 55
4.1.1 Rejection test 55
4.1.2 Pure water test 56
4.1.3 Membrane coefficient and membrane resistance 60
4.2 TDS Removal Test 61
4.2.1 TDS removal on hollow fiber sweeping gas
membrane distillation HF-SGMD
62
4.2.2 TDS removal on hollow fiber direct contact
membrane distillation HF-DCMD
66
4.2.3 Fouling analysis in MD with high concentration of salt
solution
71
4.3 Glucose Liquid Concentration Test 73
4.3.1 Glucose liquid concentration on hollow fiber direct
contact membrane distillation (HF DCMD)
74
4.3.2 Glucose liquid concentration on hollow fiber sweeping
gas membrane distillation (HF SGMD)
80
4.4 The Comparison between DCMD and SGMD 83
4.4.1 The comparison between HF SGMD and HF DCMD in
TDS removal test
83
4.4.2 The comparison between HF SGMD and HF DCMD in
glucose concentration test
84
5 Conclusions and Recommendations 87
5.1 Conclusions 87
5.2 Recommendations 88
References 89
Appendix A 96
Appendix B 99
Appendix C 119
Appendix D 124
Appendix E 126
Appendix F 128
vii
List of Table
Table Title Page
2.1 TDS Substance 4
2.2 Classification of Water 4
2.3 Maximum Allowable TDS in Boiler Water relates to Operating
Pressure
7
2.4 Comparison between NF system and RO system 9
2.5 Gelatinization Temperature of Difference Grain Starch 19
2.6 Summarize Membrane Filtration Technology 23
2.7 Summarize Four Configuration of Membrane Distillation 26
2.8 The Summarize of Membrane Characteristic Effect on Membrane
Flux
27
2.9 Dominant Mechanism Based on The Value of Membrane Pore Size
and Knudsen number
30
2.10 The Summarize of MD Operating Parameter 37
3.1 Specification of Hollow Fiber Membrane 45
3.2 Characteristic of Membrane Module 47
3.3 Methods of Analysis 52
3.4 Chemicals Used for Cleaning Membrane 53
4.1 Experimental Results of the Rejection Tests on HF SGMD 55
4.2 The Comparison Effect of Feed Inlet Temperature on Pure Water
Flux (PWF) between This Study and Other Authors.
58
4.3 The Comparison between MD and Conventional Method in
Desalination Process.
60
4.4 The Membrane Surface Temperature and TPC 61
4.5 Membrane Resistance and Membrane Coefficient Calculation Value
in SGMD Configuration
61
4.6 Summary of Na2SO4 Solution at Optimum Conditions in SGMD 63
4.7 Theoretical and Measured Concentration of Na2SO4 Solution with
HF SGMD Simulating Real Operation for Phenol Industry
Wastewater
65
4.8 Temperature Polarization Coefficient in SGMD Configuration with
High Concentration of Salt Solution
66
4.9 Membrane Resistance and Membrane Coefficient Calculation
Value in DCMD Configuration
68
4.10 Membrane Coefficient Comparison between the Study with some
Researchers in DCMD Configuration
68
4.12 Temperature Polarization Coefficient in the Test of Salt Solution
with DCMD
71
4.13 The Membrane Resistance and Boundary Layers Resistance. 71
4.14 Different Type of Resistance in MD with High Sodium Sulfate
Solution
72
4.15 The Recoverability of MD in Some Research 73
4.16 The Comparison between Evaporation and Membrane Techniques 75
4.18 Temperature Polarization in HF DCMD with Pure Water and
Synthetic Glucose Liquid
76
4.19 The Measured Real Glucose Liquid Concentration by DNS method
in DCMD configuration
79
4.20 Different Type of Resistance in DCMD with Real Glucose Liquid 79
viii
4.21 The Recoverability from Organic and Biological Fouling of MD in
Some Research
80
4.23 Different Type of Resistance in SGMD with Real Glucose Liquid 82
ix
List of Figures
Figure Title Page
2.1 Drinking water TDS standard 5
2.2 Deposits in the pipes 6
2.3 TDS removal technologies 7
2.4 Typical pressure driven membrane technology schematic 8
2.5 The difference between forward osmosis and reverse osmosis 11
2.6 ED schematic diagram 12
2.7 CDI schematic diagram 14
2.8 D – Glucose and L – Glucose shown in the linear form 16
2.9 Typical glucose syrup process 20
2.10 Four main configurations of membrane distillation 25
2.11 Transport mechanism in the pore of membrane (DCMD) 28
2.12 Concentration polarization profile in membrane distillation 31
2.13 Heat transfer in direct contact membrane distillation 33
2.14 Temperature polarization profile in membrane distillation 34
3.1 System calibration 41
3.2 Simple cross flow operation of membrane distillation 42
3.3 Experimental details 43
3.4 Experiment set up of lab scale hollow fiber direct contact membrane
distillation.
44
3.5 Experiment setup of lab scale hollow fiber sweeping gas membrane
distillation
44
3.6 Contact angle of membrane 45
3.7 Operation mechanism of direct contact membrane distillation 46
3.8 Operation mechanism of sweeping gas membrane distillation 46
3.9 Hollow fiber membrane distillation 47
3.10 Membrane Cleaning Process 54
4.1 The rejection result for HF SGMD with gas flow rate 16.6 L/min 56
4.2 The rejection result for HG SGMD with gas flow rate 25.5 L/min 56
4.3 Membrane flux variation at different temperature and gas flow rate 57
4.4 Energy consumption variations at different temperature and gas flow
rate
58
4.5 Energy ratio variations at different temperature and gas flow rate 59
4.6 Flux and concentration with high concentration Na2SO4 solution in
HF SGMD
62
4.7 Solubility of sodium sulfate vs. temperature 63
4.8 Energy consumption ratio of testing the capacity of membrane with
high concentration Na2SO4 solution in HF SGMD
64
4.9 Flux and concentration with high concentration Na2SO4 solution in
HF SGMD during simulated real operation
65
4.10 The energy consumption ratio of membrane with high concentration
salt solution in real operation in HF SGMD
66
4.11 Performance of permeate pump in HF DCMD system at different
flow rate
67
4.12 The performance of HF DCMD in the operation with high
concentration Na2SO4 solution
69
4.13 Measured concentration of Na2SO4 solution for HF DCMD in the 70
x
operation with high concentration Na2SO4 solution
4.14 Energy ratio with increasing salt concentration in HF DCMD 70
4.15 The ratio of membrane resistance and boundary layers resistance in
MD with high salt concentration solution
72
4.16 Different type of resistances in MD with high salt concentration 73
4.17 Permeate flux vs. feed synthetic glucose liquid in HF DCMD system 75
4.18 Measured synthetic glucose liquid concentration by DNS method in
DCMD configuration
76
4.19 Ratio of membrane resistance and boundary layers resistance in MD
with synthetic glucose liquid
77
4.20 Specific energy consumption in HF DCMD system with synthetic
glucose liquid
77
4.21 Permeate flux vs. feed real glucose liquid in HF DCMD system 78
4.22 Specific energy consumption in HF DCMD system with real glucose
liquid
78
4.23 Measured real glucose liquid concentration by DNS method in
DCMD configuration
79
4.24 Different types of resistances in DCMD with real glucose liquid 80
4.25 Permeate flux vs. feed real glucose liquid in HF SGMD system 81
4.26 Specific energy consumption in HF SGMD system with real glucose
liquid
81
4.27 Measured real glucose liquid concentration by DNS method in
SGMD configuration
82
4.28 Different types of resistances in SGMD with real glucose liquid 83
4.29 The flux comparison between HF SGMD and HF DCMD 83
4.30 The energy ratio consumption comparison between HF SGMD and
HF DCMD
84
4.31 The flux comparison between HF SGMD and HF DCMD in glucose
concentration test
85
4.32 The specific energy consumption comparison between HF SGMD
and HF DCMD in glucose concentration test
86
4.33 The resistance comparison between HF SGMD and HF DCMD in
glucose concentration test
86
xi
List of Abbreviations
∆𝐻𝑣 Latent heat for evaporation
µ Viscosity
AGMD Air gap membrane distillation
aw Water activity
CA Contact angle
COD Chemical oxygen demand
cp Specific heat capacity
CPC Concentration polarization coefficient
D Diffusion coefficient
DCMD Direct contact membrane distillation
Dh Hydraulic diameter
dp Membrane pore size
EPA Environmental Protection Agency
FS Flat sheet
h Heat transfer coefficient
HF Hollow fiber
Jw Permeate flux
k Fluid thermal conductivity
kb Boltzmann constant
Kn Knudsen number
LEP Liquid entry pressure
MD Membrane distillation
MF Micro filtration
Nu Nusselt number
P Total pressure
Pa Air pressure
PI Process Intensification
Pm Mean pressure within membrane pore
Pm Collision diameter of water molecule
Pr Prandlt number
PTFE Polytetrafluoroethylene
pw Vapor pressure
RO Reverse osmosis
SGMD Sweeping gas membrane distillation
TDS Total dissolved solid
TPC Temperature polarization coefficient
VMD Vacuum membrane distillation
WHO World Health Organization
휀 Membrane porosity
𝜆 Mean free path
1
Chapter 1
Introduction
1.1 Background
Industrial development in a sustainable way is seen as a challenge to humanity. Process
intensification (PI) is offered like a remarkable solution. The approach brings specific
benefits in industrial production (chemical and food industry) as well as in environmental
activities such as dramatically reducing equipment size, saving energy consumption,
increasing safety and minimizing the effect on environment (Drioli et al., 2011). The
process has been studied in many fields of food industry, in which, concentrating glucose
is a typical example. Glucose is also known as a monosaccharide (or simple sugar).
Glucose is in solid form, crystalline, colorless and very soluble in water. It has a sweet
taste, but glucose is less sweet than sucrose and fructose is sweeter than sucrose. Glucose
is sweeter by 0.6 times than sugar cane. The Glucose molecular formula is C6H12O6 [or H-
(C=O)-(CHOH)5-H)]. The boiling point of glucose is not documented because of the fact
that glucose heated to certain temperatures, hence the phase change from solid to melting
before boiling. Glucose crystals decompose in a process referred to as apparent melting. It
has a melting temperature at 146 0C, while the boiling point of water is at 100 0C. In
biology, glucose is considered the most important carbohydrate, in humans and animals,
glucose is the fixed component of blood, and it is easily absorbed by the human body.
Glucose is consumed by the cell as an intermediate metabolite and energy source. During
the start of respiration and photosynthesis of cellular in eukaryotes as well as prokaryotes,
glucose was generated as a main product. One of the extremely pure forms of glucose is
liquid glucose, which has a concentrated flavor. Liquid glucose is made from refined starch
by acid hydrolysis or enzyme treatment following the process of refining and
concentration. In which the concentration process is not only considered as an important
step to determine the quality of product but also reduction in packing, transport and
storage. Traditional concentration methods are using heat to evaporate solvents or
extraction. However, these methods can reduce the quality of the product due to take place
at high temperatures, using more energy or high investment costs and low efficiency.
Total Dissolved Solids (TDS) are the total number of charge ions, including minerals, salts
or metals that exist in a certain volume of water, usually expressed in units of mg/L or ppm
(parts per million). TDS can arise from two main sources which are natural as leaves, silt,
plankton, rocks and man-made sources such as agricultural activities, industrial production,
and urban run-off. TDS is commonly used as the primary basis for determining the level of
clean or pure water. Under the current provisions of the World Health Organization
(WHO) and Environmental Protection Agency (EPA), the TDS concentration does not
exceed 500 mg/L TDS for pure water (EPA, 1992) and not more than 1000 mg/L for
drinking water (WHO, 1996). A number of applications in the electronics manufacturing
industry require water free from TDS. Although TDS is generally not mentioned as a
primary pollutant, it does not directly relevant to human health, but TDS is used as an
indication of aesthetic characteristics of drinking water and it is also the overall index of
the presence of a variety of chemical pollutants. In addition, high TDS concentration also
significantly affects the aquatic ecosystem, industrial production and irrigation practices.
The removal of TDS from wastewater is relatively difficult. It requires advanced
technology with high investment and operation cost. The commonly current technology
2
can be divided into three main groups as membrane technologies, ion exchange
technologies and thermal technologies.
Membrane distillation (MD) is one of the emerging non-isothermal membrane separation
processes (M. Khayet and Matsuura, 2011) in which only volatile molecules can pass
through a porous hydrophobic membrane and non-vapor solute will retain in the feed side.
This separation process is based on the principle of the vapor pressure difference between
feed side and permeate side. The pressure of volatile solvent is raised by increasing the
temperature of feed solution. In contrast the temperature in permeate side would be kept
lower than feed side to decreasing the vapor pressure. The distinction of vapor pressure
between two sides of membrane leads to evaporating volatile solvent at the liquid-vapor
layer. The emphasized concept is that hydrophobic (non-wetting) membrane just allows
vapor pass through, water is retained. Therefore, volatile solvent can travel through the
membrane and gets condensed at the vapor-liquid layer.
Membrane Distillation has superior features compared to conventional membrane as using
less energy, system be able to work at low temperatures, no need to boil the feed solution
to the boiling point. Consequently, membrane distillation can be widely applied in many
different fields such as desalination, process wastewater treatment, water purification,
process intensification.
Due to the ability of intensifying concentration valuable compound of MD, the melting
point of glucose is higher than water and the dehydration process does not change the
physical properties of glucose so we can use MD to concentrate glucose liquid to higher
concentration. Similarly, membrane distillation is a potential technology to apply in
process wastewater treatment to remove TDS because of non-volatile nature and the need
of remove TDS from wastewater.
1.2 Objectives of the Study
The study target is to investigate the potential of MD technology for process intensification
application with respect to potential of cleaner production and removal of TDS from the
waste stream. To achieve the goal, this study was carried out with the following objectives:
Evaluate the process intensification ability in MD by optimizing the performance
of two configurations of membrane distillation (direct contact membrane
distillation and sweeping gas membrane distillation);
Evaluate the capability of removing TDS by MD utilizing two configurations
(direct contact membrane distillation and sweeping gas membrane distillation);
Performance evaluation of each configuration suitable for process intensification
and TDS removal by theory and experimental results.
1.3 Scope of the Study
In this thesis, PTFE membrane was used for the experimental work. For both direct contact
and sweeping gas membrane distillation, a hollow fiber membrane with a pore size of 0.45
µm size was used. Synthetic solutions/wastewater was used in the research to simulate
both glucose and TDS solutions. This study was conducted with the lab-scale and
researches were set as follows:
3
Finding the best conditions for increasing glucose concentration as well as
increasing permeate flux without TDS through the membrane by adjusting the
temperature, concentration of feeding solution or air pressure. The fouling and
energy consumption were included in this study;
Comparing the effectiveness of using these configurations in intensification
application and wastewater treatment application with energy consumption;
4
Chapter 2
Literature Review
2.1 Total Dissolve Solid
Solids in wastewater include suspended solids, settled solids, the colloidal particles and
dissolved solids. Total solid (TS) in wastewater is the remaining part after the complete
evaporation of wastewater at a temperature of 105oC. Total solids are expressed in units of
mg/L. Total solids can be divided into two components: suspended solids (can be filtered)
and dissolved solids (cannot filtered). Industrial wastewaters have high concentrations of
total dissolved solids (TDS) such as metallurgical industry; textile industry is among the
largest challenges when dealing with the wastewater remediation process.
2.1.1 Definition and its source
2.1.1.1 Definition
Total dissolved solids are substances that cannot be removed by conventional filtration
methods. TDS consists of dissolved organic and inorganic substances that contained in a
liquid in a molecular, ionized or micro-granular (colloidal soil) suspended form as prented
in Table 2.1.
Table 2.1 TDS Substance
Organic Pollutants, hydrocarbons, herbicide and soil organic matter
(humic/fulvic)
Inorganic salts
Anions Carbonates, nitrates, bicarbonates, chlorides and sulfates.
Cations Calcium, magnesium, potassium and sodium.
TDS is often taken as the basis for determining the level of clean water. According to
Water Quality Association, the different water sources can be classified based on the
concentration of TDS as Table 2.2:
Table 2.2 Classification of Water
Water system TDS concentration (ppm)
Fresh water <1,000
Brackish 1,000 – 5,000
Highly Brackish 5,000 – 15,000
Saline 15,000 – 30,000
Sea water 30,000 – 40,000
Brine 40,000 – 300,000
Wastewater from Industry 1,000 – 100,000
TDS at higher levels does not mean that it negatively affects human health. In fact, mineral
water contains high concentrations of dissolved solids, but these compounds beneficial to
human health. Environmental Protection Agency in the United States has responsible for
drinking water standard specified that TDS is a voluntary guideline.
5
Figure 2.1 Drinking water TDS standard (EPA, 1992)
2.1.1.2 TDS source
Natural sources: Certain natural sources of total dissolved solids arise from the dissolution
of rocks and soils as mineral springs, salt deposits, sea water intrusion and from the
weathering as carbonate deposits, storm water.
Man-made sources: sewage, industrial wastewater, chemicals used in the water treatment
process, point/non-point wastewater discharges, from surface run-off like urban run-off,
salts used for road de-icing, anti-skid materials, agricultural runoff and it also comes from
microbial contaminants such as viruses, bacteria from the sewage system, septic system.
2.1.2 The related environmental problem
2.1.2.1 Potential of health effects
TDS biased assessment of drinking water aesthetics rather than a health hazard. An
elevated TDS indicates as the following:
The scale can be formed because the presence of the dissolved ions may cause the water to
be salty, corrosive or brackish taste and interfere and decrease efficiency of hot water
heaters; and
A number of ions present at high concentrations exceeds primary and secondary standard
for drinking water, such as an elevated level of arsenic, nitrate, aluminum, lead, copper, etc
2.1.2.2 Effect of TDS on aquatic ecosystems
A certain level of TDS ions is necessary for aquatic life. However, high TDS
concentrations can affect the temperature and pH of the water. High TDS levels will lead
to increased water turbidity which interferes with photosynthesis and result in increased
water temperatures. Fluctuations in TDS concentrations, however, can be harmful as TDS
levels affect the flow of water into and out of an organism’s cells. Unsafe levels of TDS
can degrade and diminish aquatic life. Most aquatic ecosystems involving mixed fish fauna
can tolerate TDS levels of 1000 mg/L (Kaur, 2008). Road de-icing runoff may create saline
layers in receiving water body. These saline layers leading to reduced dissolves oxygen
levels in the hypolimnion because it does not mix with existing body water.
6
2.1.2.3 Industrial considerations
Using water containing high levels of TDS (above 500 mg/L) in industrial production can
cause negative effects. For example, in manufacturing activities such as cooling, boiler
feed water, if TDS levels are not completely removed, it will cause the hard water leading
to encrustation formation as showed in Figure 2.2.
Figure 2.2 Deposits in the pipes
2.1.2.4 Irrigation effect
The salinity of the irrigation water causes damage to crops. The plant can be dehydrated by
reversing the osmotic condition because of high salt concentrations (water will flow out of
the plant in an attempt to achieve equilibrium).
2.1.3 TDS in industrial wastewater, its measurement
2.1.3.1 TDS in industrial wastewater
Depending on the different stages in various industries, the requirement of TDS
concentration in the feeding water is different. Some industries have strict requirements for
TDS concentration as for the Pulp and Paper industry (light paper), maximum TDS
concentration of water use is 0.05mg/L, water use in Clear plastics has a TDS
concentration requirement is 200 mg/L, these conditions even lower than standard for
drinking water.
The ABMA - American Boiler Manufacturers Association also provides completely
different TDS concentrations in for each different operating conditions of steam boiler in
the table below:
7
Table 2.3 Maximum Allowable TDS in Boiler Water relates to Operating Pressure
Boiler operating pressure (bar) Total dissolve solid (ppm)
0 – 3.5 2500
3.5 – 20 3500
20 - 30 3000
30 – 40 2500
40 – 50 1000
50 – 60 750
60 – 70 625
2.1.3.2 TDS measurement
Currently there are two main ways to measure the TDS concentration that is Electrical
Conductivity and Gravimetry.
Gravimetry is more accurate. Concentration of Total Dissolved Solids is the dry weight of
the liquid through the filter when 1 liter of sample water filtered through the filter funnel
with glass fiber and then dried at 105oC until constant mass. The unit is mg/L. (Accurate to
1 part 10000 g).
Electrical Conductivity (based on conductivity correlates with the concentration of ionized
solids in water). This is a quick and affordable method although less accurate results.
2.1.4 Current TDS removal method
Removing TDS from water is essential because of high TDS concentrations affect the
surrounding environment and the need to use purified water for industrial activities,
removing TDS from water is essential. However, the removal of TDS is not easy, it needs
the technology with high investment and operating cost. Currently, several technologies
are being applied to remove TDS are presented in Figure 2.3:
Figure 2.3 TDS removal technologies
8
2.1.4.1 Membrane separation
Pressure driven membrane technology:
Operating mechanism of pressure-driven membrane processes (nanofiltration, forward
osmosis), a pressure exerted on the solution serves as a driving force to separate the
solution into permeate (pure water) and concentrate (concentrated solution) as presented in
Figure 2.4. Membrane between permeate and feeding side can be mineral, polymeric,
ceramic, or metallic. The filtration techniques will be differing in pore size, from dense to
porous membranes. Salts, macromolecules, small organic molecules, or particles can be
retained, and the applied pressure will differ depending on the type of technique (Van Der
Bruggen et al., 2003). TDS concentrations ranging from 500 to 40,000 mg/L in saline
streams can be typically removed by pressure driven membrane processes (Drewes, 2009).
PermeateFeed
Concentrate
RO
Figure 2.4 Typical pressure driven membrane technology schematic
Reverse osmosis
Reverse osmosis has a mechanism in contrast with usual osmosis. Earth's gravity creates
the permeation of water molecules through the capillaries of the filter (such as a ceramic
filter). RO operating bases on the mechanism of movement of water particles by
compressing high-pressure pumps to create a strong flow pushes the chemical composition,
the metal impurities move with high velocity, thrown into areas that has low pressure or
swept downstream and out along the exhaust (like the working principles of the human
kidney). Meanwhile, the water molecules pass through the pore filter size 0.001
micrometer size by excessive pressure, with the size of this pore filter, most of the metal
component chemicals, bacteria cannot pass through.
RO can treat water with concentration of TDS up to 40,000 – 45,000 mg/L but it cannot
completely treat non-ionized materials such as large organic molecules or gases that will
not pass through the membrane (ALLConsulting, 2014).
Advantages and drawbacks of RO filtration.
Advantages
Relatively simple, low maintenance system
The ability to remove variety small impurities as bacteria, ions, solids, dissolved
solids, viruses.
9
Disadvantages
The production of pure water yields is proportional to the temperature of the
water;
Able to leak some small amounts of singly charged ions (K+, Na+, Cl-);
Rapidly congestion membrane when water contains high suspended solids, it is
required a pre-treatment by ultra-filter membrane;
The filtration process by RO is relatively slow thus the most economical way is
that RO module will be running continuously even during non-production hours
(filtering stored treated wastewater and storing filtered water during off hours);
Chlorination chemicals should be removed by pre-treatment.
Before purchasing an RO system, it is important to implement other water saving measures
first so that the RO system is properly sized for the reduced water volumes. Otherwise, the
RO system will be under-utilized as other water saving measures is implemented.
Current technologies allow up to about 80% fresh water yields. At the optimum conditions
of temperature and minimal TDS levels, typical yields are 50%. Even with recovery rates
of 50% typical RO systems have a payback of one to two years with water cost. As an
example of a case study, an RO unit rated for around 69 m3 per day water recovery would
cost approximately $20,000 and save approximately 14.56 thousand m3 per year ($17,000
savings/year) (R.I.T, 2014).
Nano Filtration (NF)
Similar to Reverse Osmosis, diffusion is the basic mechanism in the mass transfer of
nanofiltration. Though generally, NF is cognate of membrane chemistry, the diffusion of
certain ionic solutes (such as chloride and sodium) can pass through nanofiltration
membrane, predominantly monovalent ions, as well as water. Larger ionic species,
including divalent and multivalent ions, and more complex molecules are highly retained.
The osmotic pressure is different between the solutions in two sides of the membrane is not
as great and this typically results in somewhat lower operating pressure of nanofiltration
compared with reverse osmosis. Therefore monovalent ion is diffusing through the NF
membrane along with the water. In many applications, nanofiltration replaces reverse
osmosis, due to the fact that NF requires less energy comparatively. Going by its
specifications, nanofiltration lays between RO and the ultrafiltration membrane (Faridirad
et al., 2014).
Table 2.4 Comparison between NF system and RO system
Parameter NF system RO system
The pressure needed Low (50-100psi) High (150-200psi)
Flow rate High Low
Storage Tank Depends Always needed
Fouling Low High
Corrosive effluent No Yes
Membrane life 3-5 years 1-3 years
Energy saving Yes No
TDS removal 50% (10-90%) 98%
10
NF membranes has ability to reject contaminants as small as 0.001 µm so NF is able to
reject high amount of divalent ions, metals (>99% of MgSO4), and radionuclides. For
water softening applications and remove metals, NF is considered the optimal choice.
Organic compounds are also removed to varying extents with NF membranes (Bellona and
Drewes, 2005). The nominal TDS range for NF applications is between 1,000 and 35,000
mg/L. Water recovery ranges from 75-90%, however it may require extensive pre-
treatment depending on feed water quality or application of scale inhibitors.
Advantage and Drawback
Advantages
Removes lime, iron and other problem-causing elements that salt-based softeners
cannot remove;
Treated water is non-corrosive;
High flow rate, but significantly less energy usage;
Minimal space is required;
Installation of modular construction is simple, maintain and adapt if water volume
or water quality changes.
Disadvantages (Emis, 2010)
Higher energy consumption than Ultrafiltration and Microfiltration (0.3 to 1
kWh/m³);
In some heavily polluted water, Pre-treatment is needed (pre-filtration 0.1 - 20
microns).
Limited retention for salts and univalent ions.
NF is sensitive to free chlorine (lifespan of 1000 ppm). High chlorine
concentrations should be treated with an active carbon filter or a bi-sulphite
treatment.
Forward Osmosis (FO)
In FO process, water in the feed solution at a lower osmotic pressure can pass naturally
through a semi-permeable membrane to the draw solution at a higher osmotic pressure
(Zhao et al., 2014). Unlike with RO, the mass transfer driven is osmotic pressure between
the feed and the draw solution. The high hydraulic pressure is not necessary. As a low
energy consumption of emerging membrane module, FO is promising applications in
desalination, water treatment, water purification, food processing, etc.
11
Figure 2.5 The difference between forward osmosis and reverse osmosis
FO membrane has ability to operating with feed TDS concentration ranging from 500
mg/L to greater than 35,000 mg/L. FO process can reject almost dissolved substance (more
than 95% rejection of TDS) and all particulate matters (Pei Xu, 2011).
Advantages
Forward Osmosis overcomes fouling limitation;
FO can treat dirty feed streams with high concentration of suspended solids;
FO process normally occurs in nature, so that it requires less energy consumption.
Thermal Driven membrane Technology
Membrane Distillation
Membrane distillation (MD) is a thermally driven membrane separation process that
utilizes a low heat source act as the driving force to transfer mass through a hydrophobic
microporous membrane. The driving force for mass transfer is a difference between vapor
pressure of a feed solution and the permeate. The MD is a unique membrane process that
able to maintain performance of the process (i.e., mass flux and non-volatile solute
substances rejection) almost independently of feed solution TDS concentration. MD is
probably the optimal method for the production of ultrapure water at lower cost than
conventional distillation process.
Sparingly soluble salts present control the range of TDS applications. However, according
to recent studies, scaling is not a key problem compared with other membrane
technologies. Feed solution with TDS concentration of 500 mg/L to more than 50,000
mg/L is possible. The studies also demonstrated that greater than 70,000 mg/L feed
streams can be processed. In theory MD has the ability to remove 100% of non-volatile
compounds.
The quality of product water from the MD process is equal to that of distilled water from
thermally driven processes (TDS from 2 to 10 mg/L). All solutes with higher volatility
than water (such as ammonia) will preferentially diffuse into the product water.
Advantage and Disadvantage (Lawson and Lloyd, 1997); (Liu and Martin, 2005)
Advantages
12
Theoretically, 100% rejection of colloids, ions, cells, macromolecules, and other
non-volatiles components;
Operating at lower temperatures than conventional distillation;
The operating pressure is lower than other pressure-driven membrane separation
processes;
Less affected by the removal of various substances in different conditions (e.g. pH
and salts);
Mechanical properties and chemical resistance ability are excellent;
Vapor spaces are reduced compared to conventional distillation processes;
Disadvantages
High intensity of energy (although the temperature of feed solution is usually low
grade);
Sensitive with surfactants;
Pretreatment methods are required to remove unwanted volatiles as carbonate or
ammonia. (pH control, degassing, etc.)
Electrically Driven Processes
Electrodialysis (ED)
Electrodialysis is an electrochemical separation process in which ions move through the
ion exchange membrane from a region of lower concentration to higher concentration
under the effect of electric current that is showed in Figure 2.6.
To control the motion of the ions in solution and the electrode area, the ion exchange
membrane is equipped. Normally we use two types of ion exchange membranes:
An anion exchange membrane, which allows anions (negatively charged ions) passes
through the membrane. This membrane is conductivity and waterproof even when applied
with pressure.
A cation exchange membrane, which allows anions (negatively charged ions) passes
through the membrane. This membrane is also conductivity and waterproof even when
applied with pressure.
Figure 2.6 ED schematic diagram
13
Depending on the number of stages present in ED unit, approximately 25% to 60% TDS
can be removed by this treatment process.
The Electrodialysis treatment process is also applicable for removing and/or reducing
barium, aluminum, cadmium, bromide, calcium, chloride, 90 to 95% of cyanide, 97-98%
of copper, 94 to 97% of potassium, and VOCs. HEED - High Efficiency Electrodialysis is
able to purifying water to contain less than 2 ppm TDS. This process generates the waste
less than 2% of the original contaminated water. To achieve this purity level, the
concentration of TDS in the original stream waste must not exceed 22,000 ppm. The High
Efficiency Electro-Pressure Membrane (HEEL) is an improved technology of the HEED. It
includes pre-filters and RO systems before HEED technology. According to vendor
information, it has ability to handling TDS levels up to 50,000 ppm and with water
recovery up to 99% of the feed water.
Advantages
Highly efficient process, the cation and anion exchange membranes with selectivity at least
90%, are used in many important industries.
Disadvantages
Ability of ED in treating TDS concentration is limited from 4,000 mg/L to 15,000
mg/L, although recently, ED technology reported the ability to treat high TDS
concentrations (approximately 35,000 mg/L TDS with 75 % recovery);
Pre-treatment is very necessary for Electrodialysis to control the build-up of
magnesium hydroxide, calcium carbonate (CaCO3) and iron ions;
ED is not effective at treating colloidal material, bacteria, boron or silica;
Scale formation in short time, complicated operation.
Ion exchange
Ion exchange processes based on the chemical interactions between the ions in the liquid
phase and solid phase. It is a process consisting of reversible chemical reactions between
the ions in the liquid phase and ions in the solid phase (the exchange resin). This process
depends on the type of resin ion exchange and the various types of ions. The IX resins
adsorption capacity is exhausted when the target ion reaches a prescribed breakthrough
concentration in the IX product water. To achieve high purity water quality, many
conventional IX processes are operated with mixed beds to achieve removal of both
cations and anions.
The ion exchange method is widely used in the wastewater treatment process as well as
water supply. In water treatment, the ion exchange method is commonly used to remove
the salt, reducing hardness, demineralized, nitrate reduction, color removal, removal of
metal and heavy metal ions and other metal ions in water. In wastewater treatment, ion-
exchange method is used to remove metal (zinc, copper, chromium, nickel, lead, mercury,
cadmium, vanadium, and manganese), the arsenic compounds, phosphorus, cyanide and
radioactive materials. This method allows recovery of valuable substances with a highly
water purification.
14
Advantages
The advantage of this method is very thorough and selective handling.
Disadvantages
The main disadvantage of this method is that the investment costs and operating high so
rarely used for large buildings and is often used in cases requiring high processing quality.
Capacitive deionization
Capacity deionization is one of the desalination technologies. The main operation
mechanism is removing ions by the porous electrodes. The surface of membrane porous is
applied a low voltage electric. The negative ions such as chloride, nitrate, and silica are
absorbed in the positive electrode side. Conversely the negative electrode side will attract
positive ions such as calcium, magnesium, and sodium. The difference compared with ion
exchange is the additional chemicals do not necessary. In ion exchange process, the
chemical is added to regenerate the electro sorbent, but in capacity deionization process,
the electro sorbent is regenerated by eliminating the electric field. CDI schematic diagram
is showed in Figure 2.7.
.
Figure 2.7 CDI schematic diagram
Advantages
Cost competitive for treating water with TDS <3,000 mg/L;
Very low infrastructure requirement. Compact size and mobility;
Low skill required to operate and maintenance.
Disadvantages
CDI has poor removal of uncharged substances such as organics and boron.
15
2.1.4.3 Thermal Technology
Thermal separation technologies include multi stage flash (MSF), vapor compression
distillation (VCD) and multiple effect distillation (MED) that are used for desalination.
In MSF, the feeding solution is heated until reaches operating temperature around 70 – 90 0C, the pressure is lowered, and the water "flashes" into steam. This process constitutes one
stage of a number of stages in series, normally about 15 to 25 stages in a seri, each
operating a lower temperature and pressure.
In MED, feeding solution will be passed through series of evaporators. This method can be
classified into two types MED at low temperatures (MED-LT) and MED at high
temperatures (MED-HT). For MED-LT system, the operation temperature is about 60 – 70 0C and the temperature in the last stage at around 40 0C, this process uses energy more
efficiently and longer working time than MSF system. MED-HT system uses gas stream at
higher temperature, the scale formation is controlled during operation. MED-HT system is
more widely used than the MED-LT system because it can double the performance.
VCD process includes evaporation of feeding solution, compressing vapor and finally
recovering the temperature of condensation for next batch.
Some distillation process can be combined as the hybrid technology, such as multi stage
flash and vapor compression distillation. The product of the combined process is a solution
with low salt concentration or without salt. The hybrid thermal technology can be used as
zero liquid discharge method. This system requires low energy consumption, low
investment cost, however the feeding solution need to be pretreated. Thus the sea water has
TDS concentration above 47,000 mg/L cannot be applied hybrid thermal treatment
technology.
In addition to the distillation technology, the thermal separation technologies as
evaporation and freeze-thaw have been developed for removing TDS from wastewater.
2.2 Glucose
Glucose is first extracted from dried grapes in 1747 by Andreas Marggraf. Glucose' name
comes from a word in Greek "glycos", which means sugar or sweet. The structure of
glucose was discovered in the period from the late 19th to early 20th century.
2.2.1 Definition
Glucose is a monosaccharide or in other words is simple sugar, is the most important
carbohydrate in biology. It is used by the cells as source of energy and metabolic
intermediate. Glucose is one of the major products of photosynthesis process.
Glucose also referred to as an aldohexose because it contains six carbon atoms and an
aldehyde group. Depending on the position of the OH group compared to the aldehyde
group, there are two kind of molecular formula of glucose is D-glucose and L-Glucose as
seen in Figure 2.8. D-glucose has biologically active, this form is often referred to as
dextrose (dextrose monohydrate), especially in the food industry. Another form is L-
glucose, in contrast with D-glucose; the cells cannot use it as energy.
16
Figure 2.8 D – Glucose and L – Glucose shown in the linear form
2.2.2 Physical and chemical properties
2.2.2.1 Physical properties of glucose
Glucose is in solid form, crystalline, colorless, with a melting temperature of 146 0C,
soluble in water. Density of Glucose is 1.54 g/cm3; it has a sweet taste, but not sweetener
than cane sugar (saccarose, C12H22O11). Glucose is by 0.6 times sweeter than sugar cane (if
the sweetness of sugar is 1, the sweetness of glucose by 0.6). Glucose is in the human body
as well as animals. Glucose in blood is at about 0.1% (by volume). In honey, glucose
concentration is about 30%.
2.2.2.2 Chemical properties of glucose
Glucose includes many – OH groups and CH = O group thus it has the nature of an
alcohols and aldehydes.
The nature of multifunction alcohol that is soluble precipitation of copper hydroxide
Cu(OH)2 formed blue solution.
C6H12O6 + Cu(OH)2 (C6H1106)2Cu + 2H2O
Nature of aldehydes: silver mirror reaction when glucose reacts with the AgNO3/ NH3
solution creating silver precipitate (so also called silvered).
CH2OH(CHOH)4CHO + AgNO3 + NH3 + H2O CH2OH(CHOH)4COONH4 + 2Ag +
2NH4NO3.
Precipitation reaction with Cu(OH)2: Glucose reacts with Cu(OH)2 and NaOH catalyst
forms brick red precipitate Cu2O.
CH2OH(CHOH)4CHO + 2Cu(OH)2 + NaOH CH2OH(CHOH)4COONa + Cu2O + H2O
Hydrogenation reaction: Hydrogen is added in the CH = O group creates CH2 – OH
group.
CH2OH(CHOH)4CHO + H2 CH2OH(CHOH)4CH2OH
17
Ester is formed in the reaction between glucose and anhydride acetic.
CH2OH(CHOH)4CHO + (CH3CO)2O CH2OCOCH3(CHOCOCH3)4CHO
Alcoholic fermentation reaction: This reaction creates ethyl alcohol (C2H5OH) and
carbon dioxide gas (CO2)
C6H12O6 Alcoholic fermentation C2H5OH + 2CO2
Lactic acid fermentation reaction: Glucose can be modulated ethyl alcohol with the help
of yeast is tasked as the catalyst.
OH
C6H12O6 2CH3 – CH – COOH
2.2.2.3 Glucose source
In nature
Most of the glucose is in the body plant such as roots, flowers, leaves etc. Ripe fruit
(especially grapes) is the major source of glucose.
In commercial
In the industrial scale, glucose is formed by enzymatic hydrolysis process. Glucose
production processes can use many different raw materials such as corn, wheat, rice,
tapioca, arrowroot plant depending on the climate characteristics of different area.
Cornstarch is the most common raw material in the production process glucose in the
United States.
Enzyme hydration process is divided into two main phases. Starch is hydrolyzed into
smaller carbohydrates in about 1-2 hours, at around 100oC. The heating time can be
reduced by increasing the temperature of the mixture starch up to 130oC. High
temperatures make the starch dissolves readily in water, however it also deactivates
enzyme, so fresh enzyme must be added to the mixture after each heating. The second step
in the formation of glucose process is "saccharification". Enzyme glucoamylase from
Aspergillus Niger fungus can almost completely hydrolyze starch. The best condition for
reaction is at pH 4.0-4.5 and temperature is 60 ° C. After 14 days, 96% starch is converted
into glucose.
2.2.3 Application of glucose
The physical, chemical and nutrition properties of glucose is applied in many food
industrial fields such as industrial fermentation (beer, alcoholic beverages, etc), bread
production, confectionery manufacturing industry, canned food, fast food and other fields,
such as chemical industry and pharmaceuticals.
Glucose is used in bread production to increase fermentation capacity, increase the
toughness of the crust, improves color, taste and structure of cake. In cake production,
glucose helps to increase the volume, structure, the balance of the cake. Glucose controls
18
the sweetness and flavor of the biscuit, it is covered up in the baking process to color and
soften the surface of the cake. Glucose also brings soft structure, wonderful sweetness and
good flow ability for ice cream and frozen desserts.
During fermentation, glucose is used as substrate, capable of supporting the fermentation
process to reduce the amount of calories and carbohydrate in the low-energy beers. In
wine, Glucose is used for increasing fermentation capacity, and increase the sweetness of
the product.
In beverages, glucose provides sweetness, osmotic pressure; it is also filler that helps
increase the taste. It controls mobility and increases storage time for powdered drinks. In
the production of candy, glucose supply sweet capital, softness and help control the
crystallization phenomena. The combination of glucose and sucrose help to increase taste,
color, gloss, increase cooling sensation in the mouth, as well as balance the sweetness,
toughness and hardness of the candy.
In canned sauce, vegetable soup, canned fruits, jams, fruit jellies, glucose is used to
provide sweetness and taste, durability and osmotic pressure, improve the structure and
quality aesthetic quality of the product. Glucose is also involved in the process of coloring
products such as sausages, peanut butter.
In the pharmaceutical industry, glucose is used as a liquid for intravenous administration,
or bounded into pill form. It is also used as raw material of the fermentation process in the
production of organic acids, vitamins, antibiotics, enzymes, amino acids, polysaccharides.
Glucose is the highest demand in fuel ethanol production.
Glucose is a particularly practical applications in the production of fruit juice, this is the
raw material is mixed into the juice to enhance the flavor and make products more
excellent quality and it also causes help lower production costs because glucose is material
that produced easily and cheaper than sucrose, also known as the diameter which is still
commonly used in everyday family as well as materials processing in some other products
in the food industry like confectionery, jam and soft drinks.
Glucose is essential for culturing microorganisms. It is easy to ferment sugar to create
alcohol, acetic acid, lactic acid, organic acids such as glutamic acid, citric acid.
2.2.4 Glucose syrup
The mixture of Glucose, a kind of sugar, and water is used to make sweet, viscous glucose
syrup. It is applied in many fields, such as food production, medicine, etc. Glucose syrup
could be extracted from many sources as corn, grape, etc. In which the most common
source has been known is starch. The general hydrolysis method could be expressed as
certain steps below, which is not depended on input materials:
Preparation
The first step in the production process of glucose syrup is material is crushed into small
pieces so that it can easily be dissolved in water. In addition, the separation of the
impurities from starch is also essential. The impurity compounds are usually fiber and
protein. Fiber cannot be dissolved in water so need to get rid of them if not the starch will
19
be difficult to become hydrated and it will also affect product quality. The presence of
protein in the starch will make the flavor and color of glucose syrup changed because the
Maillard reaction occurred. The amount of protein and fiber are removed can be used to
produce by-products such as animal feed or used in the textile industry.
Soaking
The starch need to be soaked until it become swell, this will create favorable conditions for
enzyme or acid activity. When the raw material is used is grain, the soaking solution
should add sulfur dioxide to preclude spoilage.
Gelatinization
The molecules of starch are chopped as the shorter circuit dextrin and a part of starch also
hydrolyzed by heating clean crushed material. The broken intermolecular bonds of glucose
will participate with hydrogen bonding easier. At this point the mixture becomes viscous,
consistent liquid, maximum swelling that creates favorable conditions for the complete
hydrolysis of starch.
Table 2. 5 Gelatinization Temperature of Difference Grain Starch (Jesse, 2007)
Grain Starches Temperature (0C)
Barley 52 – 59
Wheat 58 – 64
Rye 57 – 70
Maize/corn 62 – 72
Rice 68 – 77
Sorghum 68 – 77
Hydrolysis
Methods of enzymatic hydrolysis, acid hydrolysis or a combination of both methods can
produce glucose syrup. Previously, the only method is hydrolyzing corn starch with dilute
hydrochloric acid under high temperature and pressure to producing glucose syrup. Today,
the application of enzymes to hydrolyze starch is used quite extensively. First, α-amylase
enzyme is added to a mixture of starch and water to reduce the hydrolysis temperature. α-
amylase enzyme hardly exerted on intact starch which only acts on gelatinized starch and
make it diluted. At the maximum temperature is 85 0C, Bacillus Subtilis can produce α-
amylase enzyme. Bacillus Lichesuformis and Bacillus Stearothermophilus can withstand
temperatures up to 100 0C.
Clarification
After hydrolysis process, diluted glucose syrup is formed, it quickly passed through the
clarification system to remove impurities and improve the color as well as its stability.
The overall of glucose syrup production is expressed in Figure 2.9.
20
Raw material
Washing Chopping CrushingSoaking
Sulfur dioxide
GelatinizationHydrolysis
α-amylase
enzyme
Clarification
Concentration
Process
Liquid Glucose
Figure 2.9 Typical glucose syrup process
21
Evaporation
Syrup glucose solution after hydrolysis will has average dry matter concentration of 16%.
To raising the solid concentration in the solution, many technologies are applied as
evaporation, membrane, distillation, etc.
2.2.5 The current techniques for concentrating glucose
Although the price of glucose syrup increase gradually annually, the demand for this item
still high due to the high applicable value of glucose syrup in many fields such as food
processing industry, pharmaceutical industry, etc. Therefore, besides reducing production
costs, improving product quality is equally important. However, this study focuses only on
the final production process that is concentrated phase, in order to reduce production costs
and increase product quality. Currently, few commercially feasible methods include
vacuum evaporation, freeze concentration and membrane processes such as osmosis,
reverse osmosis and ultrafiltration.
2.2.5.1 Membrane technologies
Membranes technology has many configurations, including hollow fiber, spirally wound
and tubular have been used for a long time in the food industry. The technology can be
used in the production process as concentration or clarification step, as well as applied to
treat the wastewater created after production process or re-use (Inc, 2004). Membrane
technology has been clearly presented in the section 2.1.4.1. However, ultrafiltration can
cause the loss of sugars, amino acids and vitamins (Fellows, 2009).
2.2.5.2 Freeze concentration
Freeze Concentration is suitable method for concentrated fluids that their constituents
easily altered by the effects of high temperatures. The freeze concentration process is based
on the difference freezing point of liquid and water. The oldest form of freeze
concentration is very simple that a liquid barrel is left outside the cold winter until the
water in liquid become ice that cling to the inside wall of barrel. Water is frozen can be
separated out of the concentrated product.
Modern freeze concentration method includes a crystallization unit where a part of water is
frozen into solid form as ice crystals by using refrigeration system. After that, the filter,
centrifuges or Niro technology is applied to separating the ice crystals. However, this
method does not really efficient in commercial case.
Vacuum evaporation
One of the critical methods to concentrate or preserve beverage and food products is
vacuum evaporation. The operating principle of this method is that water is taken out by
evaporation, leave the liquid with higher density or in other word more concentrated.
Water evaporated by changing vacuum pressure. When the vapor pressure on the surface
of product is lower, the boiling temperature of product will be reduced. Thus creating
vacuum condition in concentration equipment will reduce the boiling temperature of the
product. In other words, the boiling temperature is adjusted by changing the vacuum level.
Many dairy and food products necessitate vacuum evaporation at high technical condition
to gain the product with suitable concentration as well as longer expiry date.
22
Table 2.6 Relations Between the Vacuum Pressure and the Boiling Temperature of
Water (VOER, 2000).
Vacuum pressure (mmHg) Boiling temperature (oC)
0 100
234 90
405 80
610 60
667.6 50
The typical vacuum evaporation machine consists three main parts. First part is heat
exchanger to transmit temperature from heat stream to food product. The second part is
separator between vapor and liquid form. The last important part is vacuum producer that
is usually operated by mechanical pump.
Base on the characteristic of product' liquid as well as desired concentration, the suitable
kind of vacuum evaporator will be chosen. A number of factors are need to be considered
in the selection vacuum evaporator is concentrate properties, the impacts of temperature
and timing, foaming and flavor or quality recovery. Cost is also a critical element for
manufacturers.
Vacuum pressure method can be applied to drying materials containing essential oils,
fragrance, pharmaceutical, food and agricultural products required low-temperature drying
to maintain quality and color, not causing destruction, denaturing agents. Vacuum pressure
method also used to dry high quality wood.
2.3 Membrane Technology
With demand for water reuse and or highly pure water for industrial applications, advanced
water treatment technologies are increasingly high levels of interest and development,
especially in membrane technology, which has emerged as a significant innovation for
treatment and reclamation, as well as a leading process in the upgrade and expansion of
wastewater treatment plants.
Basic principles of membrane filtration technology are the use of pressure to separate
water and soluble substances by using a semi-permeable membrane. The separation
mechanism is accomplished by fluid flow is moved selectively through a membrane, the
amount of components of the liquid is kept on the porous surface membrane depend on
their size. Depending on treatment demand, the appropriate membrane process is selected.
The common membrane technologies applied in wastewater treatment and water as
microfiltration (MF), Ultrafiltration (UF), Nanofiltration (NF), Reverse Osmosis (RO)
share the working mechanism only difference in the pore size membrane as showed in
Table 2.6.
23
Table 2. 6 Summarize Membrane Filtration Technology
Name Pore size
(µm)
Working
pressure
(bar)
Processing capability
Production
costs
Microfiltration
(MF) 0.1 – 1.0 1 – 8.6
Turbidity, suspended solids,
suspensions, colloids,
molecules, bacteria or
dissolved solids have larger
than pore size.
Low
Ultrafiltration
(UF) 0.1 – 0.01 4.8 – 13.8
Similar to MF, in addition,
virus, the molar mass of
small proteins, enzymes,
carbohydrates are also
retained.
Medium
Nanofiltration
(NF)
0.01 –
0.001 6.9 – 41.4
Similar to UF, in addition,
low valence molecular
salts, minerals, protein,
gelatins are retained.
High
Reverse
Osmosis (RO) < 0.001 27.6 – 68.9
Almost completely, leaving
only pure water. Very high
2.4 Membrane Distillation
Basically, membrane distillation is the process that is applied to separate or purify liquids.
The process works at the temperature that much lower than the boiling point of aqueous
solution. The hydrophobic nature of microporous membrane prevent liquid phase from
entering its pore due to surface tension force, it just accepts bulk transport of the gas phase
across the hydrophobic membrane. Gas phase also known as vapour of volatile compound
is created at the bulk feed of membrane surface by the difference in vapour pressure
between two sides of the membrane, in other word the difference temperature between feed
and permeate side. Then the vapor of volatile compound will gets condenses at lower
temperature.
2.4.1 Membrane distillation application
Membrane Distillation has been studied in many areas for various purposes from
wastewater treatment, desalination to applications in the food industry. However, we can
divide the application of MD into two main parts as Intensification and Purification.
2.4.1.1 Intensification
Process Intensification in MD application is the process of increasing the quantity of
compounds in a certain solution by splitting of a portion of the solvent. The process is
applied in the field of environment with the purpose of reducing transportation cost,
treatment cost, easier handling than diluted stream as applied in the food industry, for
example the study on MD in concentration orange juice (Vincenza Calabro, 1994) and in
chemical industry.
24
2.4.1.2 Purification
Purification in MD is the process of removing non-volatile undesirable substances out of
volatile compound. In MD, instead of the desired product is compounds at the feed side in
intensification process, on the other hand, in the purification process, the desired
compound is at the permeate side. This process also has been studied in various fields in
which the main application is producing fresh water from sea water (Carlsson, 1983), in
addition, the producing pure water from textile waste water (Calabrò et al., 1991).
2.4.2 Membrane distillation configuration
Membrane Distillation has four configurations that including direct contact membrane
distillation (DCMD), sweeping gas membrane distillation (SGMD), air gap membrane
distillation (AGMD), and vacuum membrane distillation (VMD) as Figure 2.10. The
fundamental distinction between four configurations is the nature of the colder side
processing of the permeate side.
Direct contact membrane distillation (DCMD)
The simplest configuration of MD is Direct Contact Membrane Distillation. Both warm
feed and cold permeate aqueous solution are contacted in direct with the hydrophobic
membrane. It has been studied for application in many fields as Fruit Juice Concentration
(Tajane, 2010), Desalination (Cheng et al., 2008) or acids manufacturing (Tomaszewska et
al., 1995). However the biggest achilles of DCMD systems is that the cool aqueous
solution on the permeate side results in large conductive heat loss through the thin
membrane (Summers et al., 2012).
Air gap membrane distillation (AGMD)
To overcome the conductive heat losses in DCMD, a stagnant air gap is added in the space
between membrane and condense surface. The obligation of this gas layer is to transport
vapor from permeate membrane surface to condenser. AGMD can be used in most
applications of MD. The configuration has the highest energy efficiency. However the
main disadvantage of AGMD is the air gap will creates additional resistance for mass
transfer leading to reduce the flux of membrane (Yarlagadda et al., 2010).
Vacuum membrane distillation (VMD)
The drawback of AGM can be eliminated by setting a vacuum pump to create negative
pressure in permeate membrane side. This strategy enhances the difference of pressure
between two sides of membrane leading to increasing the mass flux. The heat loss in this
configuration is also negligible (Alkhudhiri et al., 2012). The vapor after taken out will be
condensed in the membrane module, or in a separate condenser.
Sweeping gas membrane distillation (SGMD)
In the sweeping gas membrane distillation (SGMD) configuration, the vapor at the
permeate membrane side is swept out by an inert gas and after that condensed by an
separate condenser (Yarlagadda et al., 2010). Similar to AGMD, the inert gas is not only
responsible as the barrier to reduce the heat loss but also enhance the mass transfer
25
coefficient. SGMD has been utilized to remove volatile substances other than water
(Garcı́a-Payo et al., 2002).Because the sweeping gas is not fixation, a large volume of inert
gas just pushed out tiny volume of permeate vapor, so the condenser should has higher
capacity.
Thermostatic sweeping gas membrane distillation is combination between AGMD and
SGMD. In this modified configuration, sweeping gas pass through the space between
membrane and cold condensation surface. This cold surface is added to reduce the
increased temperature of inert gas and the vapor also can condensed at this surface, the
remaining part of vapor can be condensed by an external condenser (Garcı́a-Payo et al.,
2002; M. Khayet, 2011).
Direct Contact Membrane
Distillation
Co
ld a
qu
eou
s so
luti
on
Hot
fee
d s
olu
tion
Air
gap
Ho
t fe
ed
so
luti
on
Con
den
ser
surf
ace
Vac
uu
m
Hot
fee
d s
olu
tion
Sw
eep
ing
Gas
Ho
t fe
ed
so
luti
on
Air Gap Membrane Distillation
Vacuum Membrane Distillation Sweeping Gas Membrane Distillation
Figure 2.10 Four main configurations of membrane distillation
26
Table 2. 7 Summarize Four Configuration of Membrane Distillation
Configuration Permeate
side
Membrane
Flux
Heat
loss
Energy
consumptn
Application ability
Intensificatn Purificatn
DCMD
Cold
aqueous
solution * *** * * ***
AGMD Air gap * * * * **
VMD Vacuum *** * ** *** *
SWMD Sweeping
gas *** * ** *** *
*Remark: *** - Good, ** - Medium, * - Poor
2.4.3 Membrane characteristics
2.4.3.1 Liquid entry pressure
Liquid entry pressure also known as wetting pressure is the limitation of the difference of
vapor pressure between two sides of membrane. If the vapor pressure difference exceeds
LEP, the hydrophobic force of membrane will be fail; the feeding aqueous solution will
penetrate through hydrophobic membrane pores.
LEP extremely depend on the pore size, membrane hydrophobicity, feed concentration and
the presence of organic solutes. The calculation of LEP is estimated by following equation
(Franken et al., 1987):
∆𝑃 = 𝑃𝑓 − 𝑃𝑝 =
−2𝐵𝛾𝑐𝑜𝑠𝜃
𝑟𝑚𝑎𝑥 (2.1)
And: 𝐿𝐸𝑃 > ∆𝑃 (2.2)
Where:
Pf and Pp are respectively the hydraulic pressure on the feed and permeate side;
B is a geometric factor (cylindrical pores has B=1);
γ is liquid surface tension: Solutions contain inorganic solutes have greater surface tension
than water surface tension (72 mN/m). In contrast, the solutions contain organic solutes
will drastic reduced the surface tension of the solution (García-Payo et al., 2000);
θ contact angle, depend on material of membrane;
rmax is the maximum pore size.
2.4.3.2 Thickness
The thickness of membrane is one of critical characteristics in the MD system. The
membrane thickness has inversely effect to the permeate flux. The permeate flux increases
as the membrane becomes thinner, by the reason of the reducing mass transfer resistance,
while heat loss increase due to heat transfer. In DCMD, the membrane thickness can be
expected in the range 30–60 μm (Laganà et al., 2000). In the case that temperature between
two sides of membrane is not significantly difference, a minimum membrane thickness can
be exceeded (Alklaibi and Lior, 2005). However if this value turn to zero, the mass flux is
also zero due to the uniformity of temperature at two sides membrane. It should be noted
that in AGMD configuration, the effect of membrane thickness on the flux is negligible
27
because the mass transfer will be predominantly resisted by the presence of stagnant air
gap. (Alkhudhiri et al., 2012).
2.4.3.3 Porosity and tortuosity
Porosity (𝜺)
Membrane porosity is defined as the ratio between volume of pores and total volume of
membrane. Membrane porosity is proportional to the permeate flux and is inversely
proportional to conductive heat loss. The following equation represents the porosity of
membrane (Smolders and Franken, 1989):
휀 = 1 −𝜌𝑚
𝜌𝑝𝑜𝑙 (2.3)
Where:
𝜌𝑚 is the density of membrane;
𝜌𝑝𝑜𝑙 is the density of the polymer material.
Tortuosity (𝝉)
Tortuosity is the sinuous in membrane pore structure. The permeate flux will be reduced as
increase tortuosity. It is also an important factor in decide mechanism of mass transport
(Srisurichan et al., 2006). Tortuosity closely relate to geometric of membrane, the
calculations are also based on two kind of membrane structure (Iversen et al., 1997):
In loose packed spheres:
𝜏 =
1
휀 (2.4)
In interstices between closed packed spheres:
𝜏 =
(2 − 휀)2
휀 (2.5)
Pore size
Micro porous membrane is usually used in membrane distillation with pore size in the
range of 0.1 - 1 𝜇𝑚. The permeate flux will be increased by rising the membrane pore size.
However, the pores are too large will make it easy to get wet hence lost the selectivity
nature of hydrophobic membrane. So each MD application has a difference optimum value
of membrane pore size, it depends on the category of feed solution (El-Bourawi et al.,
2006).
Table 2.8 The Summarize of Membrane Characteristic Effect on Membrane Flux
Membrane
Distillation
Membrane characteristic
Thickness Porosity Tortuosity Pore size
Membrane Flux IP P IP P
*Remark: P is proportional, IP is inversely proportional.
28
2.4.4 Mechanism of MD transport
2.4.4.1 Mass transfer in DCMD
In DCMD, the hydrophobic membrane directly separate hot aqueous solution and cold
permeate solution. The temperature of hot aqueous solution is lower than its boiling
temperature; it is in the range of about 30 to 90 oC. The pressure of the feed and permeate
solution is similarly atmospheric and be controlled manometers. Moreover, the
hydrophobic nature of membrane can be checked by using an aqueous salt solution as
feeding solution and then measure conductivity of the permeate solution. In all Membrane
Distillation configurations, the volume of permeate solution at a certain time is used to
determine the permeate flux. In theory the flux can be calculated by equation (Khayet and
Matsuura, 2011):
𝐽𝑤 = 𝐵𝑤∆𝑝𝑤 = 𝐵𝑤(𝑝𝑤,𝑓0 𝑎𝑤,𝑓 − 𝑝𝑤,𝑝
0 𝑎𝑤,𝑝)
= 𝐵𝑤(𝑝𝑤,𝑓0 𝛾𝑤,𝑓𝑥𝑤,𝑓 − 𝑝𝑤,𝑝
0 𝛾𝑤,𝑝𝑥𝑤,𝑝) (2.6)
With 𝑝𝑤
0 (𝑇) = [𝑒𝑥𝑝 (23.1964 −3816.44
𝑇 − 46.13)] (2.7)
Where Bw, xw, T, aw, 𝛾𝑤 are the DCMD coefficient, water mole fraction, absolute
temperature, activity, activity coefficient respectively.
In the case of feed aqueous solution contains non-volatile compounds, the vapor pressure
of the solution are calculated by the following formula:
𝑝𝑤,𝑠 = (1 − 𝑥𝑠)𝑝𝑤 (2.8)
Where: x represents the mole fraction of non-volatile solute.
In MD, the mass can be transferred through two types of condition. The first condition is
permeate flux and the second one is the mass transfer through two boundary layers.
Permeate flux
The mass transport mechanism through the membrane pores can be elucidated by dusty gas
model that includes two main mechanisms:
a. Knudsen Diffusion b. Molecular Diffusion
Figure 2.11 Transport mechanism in the pore of membrane (DCMD)
29
Knudsen diffusion ( molecule – pore wall collisions )
According to (Khayet and Matsuura, 2011), the Knudsen diffusion number can be defined
as equation:
𝐾𝑛 =
𝜆
𝑑𝑝 (2.9)
Where 𝜆, dp are mean free path and membrane pore size respectively. In case of vapor
phase is water molecule, the mean free path 𝜆w can be written as:
𝜆𝑤 =
𝑘𝐵𝑇
√2 𝜋𝑃𝑚(2.641.10−10)2
(2.10)
Where kB, T, Pm, value 2.641.10−10 are Boltzmann constant, absolute temperature, mean
pressure in membrane pore, collision diameter of water molecule respectively.
The MD coefficient for Knudsen diffusion can be defined as:
𝐵𝑤
𝐷 =2휀𝑟
3𝜏𝛿(
8𝑀𝑤
𝜋𝑅𝑇)
1/2
(2.11)
Where 휀 is porosity, r is membrane pore radius, Mw is the molecular weight of water and
R, 𝜏, 𝛿 are gas constant, tortuosity, membrane thickness respectively.
Molecular diffusion (molecule – molecule collisions)
In this case, the MD coefficient of Molecular diffusion is expressed as follows:
𝐵𝑤
𝐷 =휀
𝜏𝛿
𝑃𝐷
𝑃𝑎
𝑀𝑤
𝑅𝑇 (2.12)
Where Pa is air pressure, P is total pressure and D is water diffusion coefficient.
The composition of PD can be evaluated by an empirical equation (J. Phattaranawik et al.,
2003):
PD (Pa m2/s) = 1.895 × 10−5𝑇2.072 (2.13)
The combination of Knudsen diffusion and Molecular diffusion
The combine mechanism has the MD coefficient of Molecular diffusion evaluated by
following equation:
𝐵𝑤𝐶 = [
3𝜏𝛿
2휀𝑟(
𝜋𝑅𝑇
8𝑀𝑤)
12
+𝜏𝛿𝑃𝑎
휀𝑃𝐷
𝑅𝑇
𝑀𝑤]
−1
(2.14)
The dominant mechanism can be determined based on the value of membrane pore size
and Knudsen number as the following table:
30
Table 2.9 Dominant Mechanism Based on The Value of Membrane Pore Size and
Knudsen number
Kn <10 0.01 – 10 >10
dp > 0.01𝜆𝑤 0.01𝜆𝑤 – 1000 𝜆𝑤 < 1000 𝜆𝑤
Critical
mechanism
Knudsen diffusion Knudsen diffusion and
Molecule diffusion occur
simultaneously
Molecule diffusion
The mass transfer through two boundary layers – Concentration Polarization (CP)
In MD, feed and permeate side is separated by hydrophobic membrane so only volatile
component can go through this membrane layer. The evaporation of volatile compounds at
the bulk feed leading to rise up the concentration of nonvolatile compounds at the surface
of membrane. On the other hand, the volatile compound concentration at the membrane
surface is lower than in the bulk feed.
The concentration of solute components is equal between bulk feed and entrance surface
membrane, therefore, the impact of Concentration Polarization on flux is negligible, this
information has been proved by (Ali et al., 2013) through following equation:
𝐶𝑚𝑓
𝐶𝑏𝑓= 𝑒𝑥𝑝
𝐽𝑤
𝜌𝐾𝑐 (2.15)
Where the subscripts mf, bf refer to membrane feed and bulk feed, 𝜌, 𝐾𝑐 are the density
and mass transfer coefficient.
The resistance of CP on flux is measured by CPC that is the ratio between the
concentration of nonvolatile compounds in the feeding membrane surface and the
concentration of nonvolatile compounds in the bulk feed.
𝐶𝑃𝐶 =
𝐶𝐵𝑚
𝐶𝐵𝑏 (2.16)
After a period of operation, the concentration of non-volatile substances in the feeding
solution rises gradually, leading to the density and viscosity of the feed solution increase.
This will contribute to the influence of the Temperature polarization which will be
discussed in the next section.
31
Boundary Layer Boundary Layer
Bulk PermeateBulk Feed Membrane Module
CB,b
CA,b
CB,m
CA,m
A: Volatile compounds
B: Non-volatile compounds
Figure 2.12 Concentration polarization profile in membrane distillation
(Pal and Manna, 2010)
2.4.4.2 Mass transfer in SGMD
Due to differences in temperature or in other words the difference vapor pressure between
the feed and permeate side, the vapor can pass through the hydrophobic membrane.
Transport mechanism consists of three steps, firstly, volatile compounds start evaporating
in the hot membrane surface, secondly, vapor will be pushed through the hydrophobic
membrane by vapor pressure difference, finally, an inert cold gas sweep this vapor out of
the membrane module and get condense in external condenser.
The flux of vapor through membrane (Jw) basically depends on the net membrane
coefficient (Bw) and net difference of vapor pressure (∆pw), in other word, the flux can be
expressed by membrane coefficient that includes bulk boundary layers (B’w) and the vapor
pressure difference corresponding to the bulk phases (∆p’w). The following equation
demonstrates this relationship:
Jw = Bw × ∆pw = B’w×∆p’w (2.17)
The net pressure difference also presented as the equation:
∆pw = pw0
,f × aw,f – pw,p (2.18)
Where p, a are mention about partial pressure and activity of water respectively. The
superscript w0, f, p refer to pure water, feed and permeate respectively. The net
temperature of feed solution (Tm,f) decides the value of pw,f and aw,f , whereas the net
temperature of sweeping gas (Tm,p) determines the value of pw,p . The partial pressure of
permeate can be expressed as the following quotient:
𝑝𝑤,𝑝 =
𝜔 × 𝑃
𝜔 + 0.622 (2.19)
32
Where 𝜔 is the humidity ratio and P is total permeate pressure. Furthermore, humidity
ratio is defined by the relationship between flow rate of sweeping gas (ṁ), effective
membrane area (A), permeate vapor flux (Jw) and the inlet humidity ratio (win):
𝜔 = 𝜔𝑖𝑛 +
𝐽𝑤𝐴
ṁ𝑎 (2.20)
The synthesis equation is expressed as the following quadratic equation:
𝐽𝑤2 + 𝑏𝐽𝑤 + 𝑐 = 0 (2.21)
The coefficients b and c in the equation are:
𝑏 = (𝜔𝑖𝑛 + 0.622)
ṁ𝑎
𝐴+ 𝐵𝑤(𝑃 − 𝑝𝑤,𝑓
0 𝑎𝑤) (2.22)
𝑐 = 𝐵𝑤
ṁ𝑎
𝐴[𝑃𝜔𝑖𝑛 − 𝑝𝑤,𝑓
0 𝑎𝑤(𝜔𝑖𝑛 + 0.622)] (2.23)
In the case of the feeding solution contains at least two volatile compounds and the effects
of coupling are negligible, the total flux (J) in SGMD can be expressed as following
equation:
𝐽 = ∑ 𝐵𝑗(𝑝𝑗,𝑓0 𝑎𝑗,𝑝 − 𝑝𝑗,𝑝) = ∑ 𝐵𝑗(𝑝𝑗,𝑓
0
𝑗𝑗
𝛾𝑗,𝑓𝑥𝑗,𝑓 − 𝑓𝑗,𝑝𝑃𝑦𝑗,𝑝) (2.24)
Where x, y, 𝛾 and 𝑓 are respectively the mole fraction in liquid phase, mole fraction in the
gas phase, activity coefficient and fugacity coefficient of j. The subscript j refers to all
volatile compounds.
Similar to DCMD, the mechanism of mass transfer is also decided by dusty gas.
2.4.4.3 Heat transfer in DCMD
Similar to mass transfer, heat transfer also appeared at two boundary layers and within the
membrane module. In which, heat transfer at two boundary layers will be discussed in
Temperature Polarization part, heat transfer through membrane module (Qm) can be
expressed into two mechanisms includes heat conduction through membrane material, gas
filled in the membrane pores (Qc) and latent heat that accompanied with the vapor (Qv)
(Khayet and Matsuura, 2011):
𝑄𝑚 = 𝑄𝑐 + 𝑄𝑣 (2.25)
Following two equations are used to calculate conduction heat and latent heat that
associated with vapor respectively:
𝑄𝑐 = −𝑘𝑚
𝑑𝑇
𝑑𝑥=
𝑘𝑚
𝛿(𝑇𝑚,𝑓 − 𝑇𝑚,𝑝)
(2.26)
33
Where km is membrane thermal conductivity, 𝛿 is membrane thickness, x is the distance
between membrane surfaces. Tm,f and Tm,p respectively are temperature at the feed
membrane surface and permeate membrane surface.
𝑄𝑣 = 𝐽𝑤 × ∆𝐻𝑣,𝑤 (2.27)
Where Jw, ∆𝐻𝑣,𝑤 are water flux and latent heat of vapor molecule.
Feed Boundary Layer Membrane Module Permeate Boundary Layer
Conduction
Latent Heat
Qc
Qv
Qp
Tb,f Tm,f Tm,p Tb,p
x ThicknessThickness
Figure 2.13 Heat transfer in direct contact membrane distillation
(Khayet and Matsuura, 2011) also indicates that latent heat consumes 50 – 80 % of energy
for producing vapor while the heat used un-effectively in DCMD is thermal conduction, in
the other work it can be called as heat lost. If the operating feed temperature is higher, the
heat loss will be less considerable.
Temperature Polarization (TP)
The presence of feed boundary layer and permeate boundary layer are also responsible for
the heat transfer.
At the feed boundary layer:
𝑄𝑓 = ℎ𝑓 × (𝑇𝑏,𝑓 − 𝑇𝑚,𝑓) (2.28)
At the permeate boundary layer:
𝑄𝑝 = ℎ𝑝 × (𝑇𝑚,𝑝 − 𝑇𝑏,𝑝) (2.29)
Where hf is the coefficient of heat transfer in feed side and hp is the coefficient of heat
transfer in the permeate.
However, because of the high viscosity of feed solution and low flow rate, the temperature
in the bulk and membrane surface will not be the same, this phenomena is called
34
temperature polarization as Figure 2.14. These boundary layers impose the resistance to
heat transfer leading to reduce membrane flux. Temperature polarization coefficient (TPC)
used to measure the impact of TP on the driving force of mass transfer. TPC can be
expressed as:
𝑇𝑃𝐶 =
𝑇𝑚,𝑓 − 𝑇𝑚,𝑝
𝑇𝑏,𝑓 − 𝑇𝑏,𝑝 (2.30)
TPC value from 0.4 – 0.7 is satisfy for DCMD (Jirachote Phattaranawik and Jiraratananon,
2001), this value is lower in the membrane has poorly design. This situation can be
mitigated by creating turbulent regimes. B
oun
dar
y L
ayer
Tbf
Tm,f
Tm,p
Tb,p
Boun
dar
y L
ayer
Bulk
Per
mea
te
Bu
lk F
eed
Membrane Module
Figure 2.14 Temperature polarization profile in membrane distillation
(Pal and Manna, 2010)
The heat transfer is the same in two boundary layers and membrane module at steady state
condition:
𝑄𝑓 = 𝑄𝑝 = 𝑄𝑚 (2.31)
Or: ℎ𝑓(𝑇𝑏,𝑓 − 𝑇𝑚,𝑓) = ℎ𝑝(𝑇𝑚,𝑝 − 𝑇𝑏,𝑝) =
𝑘𝑚
𝛿(𝑇𝑚,𝑓 − 𝑇𝑚,𝑝) + 𝐽𝑤∆𝐻𝑣,𝑤
= 𝐻(𝑇𝑏,𝑓 − 𝑇𝑏.𝑝) (2.32)
Where H refers to the coefficient of global heat transfer.
The temperature at surface of membrane cannot be measured directly. It can only be
calculated by theory through heat transfer coefficient formula (Martı́nez-Dı́ez and
Vázquez-González, 1999).
ℎ =
𝑁𝑢𝑘
𝑑ℎ (2.33)
35
Where k, dh are fluid thermal conductivity and hydraulic diameter respectively.
Nu is Nusselt number (Nu) is defined as the ratio between convective and conductive heat
transfer through boundary layers that is calculated by Equation 2.34.
𝑁𝑢 = 1.86 × (𝑅𝑒 × 𝑃𝑟 ×
𝐷ℎ
𝐿)
1/3
(2.34)
Pr is Prandtl number that is evaluated by following formula:
𝑃𝑟 =
𝜇𝐶𝑝
𝑘 (2.35)
The temperature at surface of membrane can be determined by Equation 2.31 that is
obtained from Equation 2.32:
𝑇𝑚,𝑓 =
𝑘𝑚
𝛿(𝑇𝑏,𝑝 +
ℎ𝑓
ℎ𝑝𝑇𝑏,𝑓) + ℎ𝑓𝑇𝑏,𝑓 − 𝐽𝑤∆𝐻𝑣,𝑤
𝑘𝑚
𝛿+ ℎ𝑓 (1 +
𝑘𝑚
𝛿ℎ𝑝)
(2.36)
𝑇𝑚,𝑝 =
𝑘𝑚
𝛿(𝑇𝑏,𝑝 +
ℎ𝑝
ℎ𝑓𝑇𝑏,𝑝) + ℎ𝑝𝑇𝑏,𝑝 − 𝐽𝑤∆𝐻𝑣,𝑤
𝑘𝑚
𝛿+ ℎ𝑝 (1 +
𝑘𝑚
𝛿ℎ𝑓)
(2.37)
2.4.4.4 Heat transfer in SGMD
The heat transfer through membrane Qm in SGMD also divided by two mechanisms: Qc is
conduction inside membrane pores, gas-filled in membrane pores and Qv is latent heat.
However, at the same condition, thermal efficiency in SGMD is higher than its DCMD
because the heat loss by conduction is lower and decrease with increasing feeding
temperature.
The heat transfer through two boundary layers as feed boundary layers and permeate
boundary layer can be expressed as following equation:
𝑄𝑓 = ℎ𝑓 × (𝑇𝑏,𝑓 − 𝑇𝑚,𝑓) (2.38)
𝑄𝑎 = ℎ𝑎 × (𝑇𝑚,𝑝 − 𝑇𝑏,𝑝) (2.39)
Where hf and ha are heat transfer coefficient of feeding solution and sweeping gas
respectively.
hf can be calculated by Equation 2.33 to 2.35
ha can be calculated by following equation (M. Khayet et al., 2000):
ℎ𝑎 = 0.206 (
𝑘
𝑑ℎ) (𝑅𝑒. 𝑐𝑜𝑠𝛼)0.63𝑃𝑟0.36 (2.40)
36
Where k is thermal conductivity, dh is the equivalent diameter of flow channel, yaw angle
𝛼 is 0 for cross flow and the parallel flow, this value is 90.
However, in stable condition, the heat loss from sweeping gas to the environment is
negligible, the heat transfer through permeate boundary layer can be represented by the
following formula:
𝑄𝑎 =
ṁ𝑎(𝑐𝑎 + 𝜔𝑖𝑛𝑐𝑤)(𝑇𝑎,𝑜𝑢𝑡 − 𝑇𝑎,𝑖𝑛)
𝐴+ 𝐽𝑤(∆𝐻𝑣
0 + 𝑐𝑤𝑇𝑎,𝑜𝑢𝑡) (2.41)
2.4.5 Operating parameter
2.4.5.1 Temperature
Feed temperature
In MD, the feed temperature should be really lower than the boiling point of feeding
solution, normally it is in the range from 20 to 90 0C. This is the important parameter in
both DCMD and SGMD that used for controlling membrane flux. The membrane flux
exponential increase with the increasing of vapor pressure or in other word, the feeding
temperature. However the increase in the temperature polarization is also occurred when
we imply a higher feeding aqueous solution resulting in reducing the flux. Additionally, if
the increase of thermal efficiency is also taken into consideration, operation at higher feed
temperature is advisable. The influence of temperature polarization on SGMD flux is less
than its effect on DCMD flux due to air boundary layer in the permeate side of SGMD. In
SGMD, the temperature polarization coefficient in permeate side also lower than in the
feed side. The feeding temperature has negligible effect on the flux in SGMD.
Permeate temperature in DCMD
In DCMD, the permeate temperature must be lower than the feed temperature; it is in the
range of 10 – 40 0C. In contrast to the effects of feeding temperature on membrane flux, an
increasing temperature in permeate side leading to reducing the permeate flux. Because the
flux through the membrane depends critically on the temperature difference between the
two sides membranes, thus feeding temperature should be high and the permeate
temperature should be kept in minimum value.
Gas temperature in SGMD
In SGMD, the sweeping gas temperature is in the range between 100C and 300C. However,
this gas temperature increases more quickly when reduce membrane module length. The
temperature of sweeping gas increases from inlet to outlet faster than the cold solution used
in DCMD. (Basini et al., 1987) proved that despite rapidly rising temperature of sweeping
gas during operation, but it does not significantly affect the membrane flux. The influence
of air temperature to the membrane flux can be reduced by increasing the length of the
membrane module.
37
2.4.5.2 The flow rate of feed and permeate solution
In DCMD, the membrane flux increases with raising the flow rate because of the decrease
of temperature polarization. However, the varied flow rate should be encouraged to prevent
wetting the hydrophobic membrane pore, and the applied pressure also needs to be lower
than LEP value.
In SGMD, because the gas flow rate in the permeate side is governed the temperature
polarization so the permeate flux is less sensitive with the feeding flow rate. In contrast,
sweeping gas flow rate plays an important role in membrane flux. The flow rate of gas
should be varied to find out the optimum value for each membrane module. The flow rate
of sweeping gas increases that means increasing Reynolds number or other words is
changing of regime from laminar to transitional and finally turbulent regime. This regime
will reduce the effect of temperature polarization thus enhance the heat transfer coefficient
leading to increase the membrane flux.
2.4.5.3 The concentration of nonvolatile in feeding solution
The efficiency of MD in term of permeate flux reduces as the increase of concentration of
nonvolatile compounds in the feed solution (Martínez, 2004). The high feed concentration
imposes a negative effect on thermal efficiency and temperature polarization coefficient.
The rising concentration of feeding solution or in other words, that is the increase of mass
fraction ω of non-volatile substances will lead to reducing water activity (aw), and it has
been proven in the following formula (for sugar liquid at 25oC):
𝑎𝑤 = −0.27𝜔3 − 0.08𝜔2 − 0.09𝜔 + 1 (2.42)
According to the Equation (2.6), the flux through membrane will be decreased due to the
reducing of water activity. Furthermore, concentration polarization phenomena also
impose a resistant on mass transfer.
For highly viscous fluid such as sucrose liquid, high viscosity will cause a significant
impact on the heat transfer coefficient in the feeding boundary layer thus reducing the flux
through the membrane.
Table 2.10 The Summarize of MD Operating Parameter
Membrane
Flux
Operating Parameter
Feed side Permeate side
Temperature Flow rate Concentration Temperature Flow rate
DCMD P P IP IP P
SGMD P N IP N P
*Remark: P is proportional, IP is inversely proportional and N is negligible
2.4.6 Fouling and solution
2.4.6.1 Fouling
(Gryta, 2008) has demonstrated that fouling membrane modules significantly affect the
flux through the membrane. It also enhanced temperature polarization problem as a fouling
38
layer is created in front of membrane surface that imposes a heat transfer resistance. The
main reason leading to membrane is blocked is that pressure acting on the feed membrane
side is greater than the LEP value.
Based on the structure, the fouling layer created on the feeding membrane surface divided
into two basic configurations: porous and homogeneous (non-porous). The significant
factor decreases mass transfer through the membrane is caused by the formation of the
homogeneous fouling layer. The permeate flux value can be exponentially reduced by the
formation of non-porous fouling layer.
Based on the type of substance causing membrane fouling, there are three main kinds as
follow:
Crystallization fouling
The process of saturated salt solution in the feed side of the membrane has promoted the
formation of crystallization salt layers. Because of the effects of concentration polarization,
this process takes place mainly in surface of membrane module. This layer will make the
vapor pressure increased in the feed side until it exceeds the LEP value leading to wetted
membrane and loss the selectivity of hydrophobic membrane. In the desalination
application of MD, crystallization is a dominant problem.
Organic fouling
According to (Gryta, 2008), organic material is absorbed onto the feeding membrane
surface because of the hydrophobic characteristic of the membrane. Protein tends to be
absorbed by the membrane, thus formed on the membrane surface a deposit layer. So, the
author also recommends that the wastewater treatment containing proteins, polysaccharides
and amino sugars by the MD process is very complicated.
In addition, the biological microorganism like algae, fungi, bacteria and macro organism
can also stick fast to the membrane surface as well as the organic layers and rapid
developing caused the wet membrane pores and membrane flux reduction.
2.4.6.2 Solution
However, compared with other types of membrane process, fouling less severe impact on
the MD (Srisurichan et al., 2005). This problem can be overcome by applying three
common methods as membrane material modification, selection of appropriate
pretreatment or restore membrane material.
Membrane material modification
The interaction between congestor and membrane surface can be minimized by changing
the velocity as well as turbulent flow regime of feeding flow. However, over time,
congestor will in contact with the membrane. Thus, membrane fouling can be reduced by
preventing contact between the membrane and foulent. Therefore, the development of new
membrane materials and changing of membrane surface structure is necessary to reduce
fouling of membrane.
39
Appropriate pretreatment method.
Selection of pretreatment methods is very important; it can eliminate most of the
suspended solids and depending on the selected technologies that part of dissolved solids
will be removed.
Clean and restore membrane material
The easiest way to clean the membrane that is using pure water for rinsing the membrane,
hydraulic force of the water will sweep a part of foulent out of the membrane surface, the
rest is more difficult to remove and known as reversible and irreversible fouling.
However, reversible fouling also can be removed by using chemicals to clean the
membrane. The selection of chemicals must be conducted carefully because the
inappropriate chemical will not only affect the membrane surface, the fouling but also
cannot be removed. The type of foulent decides the type of chemicals need to be used to
clean the membranes. For example for inorganic contaminants, acid solution (0.1wt. %
Oxalic acid and 0.8wt. % Citric acid) should be used, and on the other hand, for foulent is
organic, alkaline solutions are selected (NaOH).
The cleaning ability of chemicals depends intimately on temperature, contact time and
concentration of chemicals. At high temperatures, high concentration and long exposure
time, the percentage of reversible membrane is increased (Madaeni and Mansourpanah,
2004).
2.5 Research Gap
TDS in Industrial wastewater is among the biggest challenges while dealing with the
wastewater remediation process. Moreover, in some process of production, water without
TDS is required. Therefore, some authors have conducted experiments with the purpose is
finding suitable method to remove TDS from water/ wastewater. The conventional method
has been studied such as membrane separation, ion exchange technology and thermal
technology. However, most of these processes consumed a lot of energy and limited by
concentration of feed solution due to fouling phenomena as well as low percentage of
removed TDS. Recently, one of critical application of membrane distillation has been
exploited that is application in wastewater treatment process. Four configurations of
membrane distillation have been considered in literature to find the TDS removal ability of
membrane distillation. The results showed that MD can remove 100% of TDS. However,
no significant work has been reported in the literature on comparison between
configurations of MD based on treatment ability, energy consumption and also fouling
phenomena.
A different application of MD is process of intensification capability. This application has
been studied by some authors in the field of water production and in food industry. The
application of MD in concentrating food liquid is a great potential as compared to the
conventional methods with many drawbacks such as working at high temperature that can
leading to product degradation, consuming high energy and low concentrating capacity.
Three of four configurations of MD have been studied in this sector as DCMD, VMD, and
AGMD. The remaining configuration has not been implemented is SGMD, especially
hollow fiber membrane. The comparison in fouling, energy consumption, and
intensification ability of four configurations of membrane distillation did not considered in
previous studies.
40
Some major points are drawn from literature review:
The negative effect of TDS on environment such as aquatic eco system, potential of
health effect, irrigation effect and industrial consideration, the conventional
methods used to remove TDS;
Glucose is applied in many food industries such as industrial fermentation, bread
production, confectionery industry, etc. The common methods utilized for
concentrating Glucose liquid have high energy consumption;
MD has two main applications as intensification and purification that are
appropriate for concentrating Glucose and removing TDS respectively. This is a
potential method because of low energy consumption and high efficiency.
Therefore, the comparison between configurations of MD, in terms of overall (flux,
energy consumption) should be considered to find out the best configuration.
41
Chapter 3
Methodology
3.1 Methodology Framework
Two major phases were comprised in this study to gain the objectives presented in
Chapter 1. The first phase focused on the intensification application of membrane
distillation, whereas, the second phase was dealt with removing TDS by membrane
distillation technology. In the first phase synthetic and real glucose syrup were chosen as
concentrated substance, TDS as Na2SO4 (divalent salts) was the substance which was
removed in the second phase. The study was conducted in hollow fiber direct contact
membrane distillation and sweeping gas membrane distillation configurations.
3.1.1 System calibration
The clean membrane distillation module was checked before conducting the experiments.
The experiments with distilled water were used to evaluate the coefficient and resistance of
hydrophobic membrane. The testing on salt solution was conducted to ensure only volatile
compounds could pass through hydrophobic membrane, non-volatile compounds were
entirely retained in the feeding side. The system verification is expressed as Figure 3.1:
Hollow fiber membrane
0.45 µm
System calibration
Pure water test
Distilled water
50
60
70
Feed temp
(0C)
2.4
2
1.5
Feed flowrate
(L/min)
25.5
22.5
16.6
Gas flow rate
(L/min)
Maximum flux
Energy consumption
Temperature polarization
Analyzed
parameter
Salt rejection
Salt solution (1%)
70
Feed temp
(0C)
2.4
Feed flowrate
(L/min)
25.5
16.6
Gas flow rate
(L/min)
Maximum flux
Energy consumption
Temperature polarization
Analyzed
parameter
Figure 3.1 System calibration
42
3.1.2 The research road map
The feasible results from calibration experiments allowed of conducting the main research
on synthetic solutions/food liquid. The systematic manner of research is shown as the
Figure 3.3.
3.2 Experimental Set up
The simplest configuration in membrane distillations is DCMD. In this configuration, the
circulating pump or agitator supported the circulatory of feed and permeate solution. In
SGMD configuration, the feed aqueous solution is also tangentially circulated to the
surface of membrane. However in the permeate side, instead of cool water, the inert air
was circulated to sweeping out the permeate flux that is condensed in condenser. The
simple Figure 3.2 illustrated to explain the cross flow mode of membrane distillation
system:
Recirculated feed in Recirculated feed out
Recirculated permeateRecirculated permeate
Figure 3.2 Simple cross flow operation of membrane distillation
The lab scale schematic of DCMD and SGMD were respectively simulated as Figure 3.4
and Figure 3.5. In the model, the heater was also served as the feed tank. The heater was
equipped with a thermal sensor that was responsible for maintaining the setting
temperature, which means that it disconnected power when the temperature reached the
setting temperature and automatically heated when the temperature dropped. It also was
covered by an insulation layer to minimize heat loss to the environment. The control box
was used to monitor the temperature of feed, permeate flux as well as cool water in DCMD
or gas in SGMD. In DCMD, the chiller and heat exchanger were provided to cooling the
water. In the other hand, in SGMD, the air compressor was supplemented to supply
sweeping gas for the model; it was measured by gas flow meter. For tracking the amount
of energy used during the initial heating of water as well as the energy required to re-heat
the water, an energy meter was installed.
43
Feed temp. : 70oC (TDS test)
– 50oC (Glucose test)
Gas flow rate : 25.5 L/min
Feed flow rate : 2.4 L/min
Time of 1 batch : 8 hrs for
TDS test – 10 hours for
Glucose test.
Feed temp. : 70oC (TDS test)
– 50oC (Glucose test)
Feed flow rate : 1.32 L/min
Permeate flow rate: 2.4 L/min
Permeate temp.: 15 oC
Time of 1 batch : 8 hrs for
TDS test – 10 hours for
Glucose test.
High concentration of divalent salt (Na2SO4)
40 g/L – 450 g/L
Analyzed parameter:
Energy consumption ratio (Kwh/ Kg)
Flux (Kg/m2.h)
Concentration (g/L)
Hollow fiber SGMD Hollow fiber DCMD
Checking permeate pump
flowrate
Hollow fiber Membrane (0.45µm)
Pure water test
Feed flow rate : 2.0; 2.4 L/min
Feed Temp : 40, 50, 60, 70 oC
Gas flow rate
Rejection test
Salt conc. : 1 g/L
Feed Temp : 70 oC
Gas flow rate: 16.6; 25.5 L/min
Best Performance Condition
- Maximum Flux;
- Minimum energy
consumption.
Removal Process
Intensification Process
Glucose liquid
35% - 60 %
Synthetic glucose
liquid
10% - 60%
Figure 3.3 Experimental details
44
The experiment setups in DCMD and SGMD were simulated by Figure 3.4 and Figure
3.5 respectively:
70
Hollow fiber
module
Feed tank = Heater
Water flow
rate meter
ValvePump
T
Chiller
T
TPermeate tank
Heat exchanger Electric meter
Figure 3.4 Experiment set up of lab scale hollow fiber direct contact membrane
distillation.
70
Hollow fiber
module
Feed tank = Heater
Electric meter
Air compressor
Air flow rate
meter
Water flow
rate meter
T
T
Valve Pump
Figure 3.5 Experiment setup of lab scale hollow fiber sweeping gas membrane
distillation
45
3.2.1 Hollow fiber membrane
The membrane was used in the study was specifically designed by a Sumitomo Electric
company with high hydrophobic ability. Moreover, the HF membrane achieved high
specific surface area, in other words, filtration area per volume of membrane is high
(m2/m3). In commercial point of view, HF is more attractive than flat sheet membrane.
However, this characteristic limited the filtration ability of membrane, especially for
wastewater contain large amount of impurity so the pre-treatment method should be
considered reasonable. In this study, only synthetic solution and pure liquid from food
process were used thus, the effect of impurity was negligible.
The mode operation of flow in hollow fiber membrane can be inside-out or outside-in. In
this study, inside-out operation was selected.
The hydrophobic characteristic of membrane depends critically on contact angle.
Normally, the material has contact angle larger than 900 is hydrophobic. The contact angle
of membrane used in the study was 112 o that is showed in Figured 3.6.
Figure 3.6 Contact angle of membrane
Details of the membrane that was used in the experiments are presented in the following
Table 3.1:
Table 3.1 Specification of Hollow Fiber Membrane
Description Characteristics
Company Name Sumitomo Electric Industries, Ltd.
Membrane Name TB – 21 – 02
Type No. 130529 – 1
Module Configuration Hollow fiber
Membrane Material Polytetrafluoethylene (PTFE)
Type of membrane Hydrophobic Microporous
Contact angle (0) 112
Nominal Pore Size (µm) 0.45
Outside Diameter (mm) 2.03
Inside Diameter (mm) 1.07
Total Length (mm) 500
Effective Length (mm) 400
Thickness (µm) 480
Number of Elements 100
Membrane Effective Area (m2) 0.255
Operating Temperature Range (0C) -100 – 260
pH Range 0 – 14
46
3.3.2 Membrane configuration
The first configuration was used in the study is Hollow Fiber Direct Contact Membrane
Distillation (HF DCMD). DCMD is the most simple among four configurations. The cool
solution in permeate side was not only assistance in creating the vapor pressure difference
between two sides of membrane, it also served as condenser for permeate flux, so the
external condenser was not required.
Cold
aqueo
us
solu
tion
Hot
feed s
olu
tion
Figure 3.7 Operation mechanism of direct contact membrane distillation
The second configuration was Hollow Fiber Sweeping Gas Membrane Distillation
(HFSGMD). In term of heat loss, this is the advanced configuration than DCMD because
the cool water was not directly contact with membrane surface, only inert gas was
provided to sweep the permeate out of membrane module. Therefore, the flux in SGMD is
higher than DCMD. However the external condenser should to be equipped in this
configuration.
Sw
eep
ing
Gas
Ho
t fe
ed
so
luti
on
Figure 3.8 Operation mechanism of sweeping gas membrane distillation
47
3.2.3 Membrane module
Unlike the flat sheet membrane, hollow fiber membrane module was not versatility and the
fabrication was relatively complex. The hollow fiber membrane was installed permanently
in the tubular module, hence, the replacing or cleaning was quite difficult. The specific
characteristic of membrane module is presented as following table:
Table 3.2 Characteristic of Membrane Module
Description Characteristic
Type of Membrane Module Hollow Fiber
Module Configuration DCMD, SGMD
Frame Material Polysulfone, PolyUrethane glue
Driving Force Thermal Driven
Inner Space (cm3) 763.72
Pipe Diameter (mm) 6
Dimension of Module (cm)
Diameter
Length
4.8
40.5
Operating Temperature (0C)
Poly sulfone: -100 to 150 oC
Polyurethane adhesive: -25 to 100 oC
-25 to 100
pH Range 2 – 13
The membrane module is expressed as Figure 3.9
Figure 3.9 Hollow fiber membrane distillation
48
3.3 Experimental Procedure
3.3.1 Flow rate calibration
Flow rate is one of important factors affect to the membrane flux. The permeate flux of
membrane increase when increasing flow rate. In addition, high flow rate can reduce the
thickness of boundary layers, consequence in minimizing the negative effect on the flux of
temperature polarization phenomena.
Beside the temperature, flow rate is also the decisive factor for the vapor pressure
difference between two sides of membrane. The higher flow rate resulted in higher of
different pressure between feed and permeate side, it also leading to increase membrane
flux.
However, the different vapor pressure between is limited by the value of LEP. If it is
higher than LEP, membrane pore is wetted. Thus, the flow rate was adjusted satisfactory, it
was not only high enough to create the turbulent condition, but also was low enough to
make the different pressure lower than LEP value of membrane. Furthermore, the flow rate
value of feed and permeate was as closer as possible.
3.3.2 Temperature calibration
Temperature in the study was observed by thermal meter in feeding point, permeate point
and also gas producer point.
The result of increase temperature of feed aqueous solution is increasing membrane flux.
However, the maximum allowable temperature of feed solution in membrane distillation is
only 90 0C. In addition, heating to high temperature consumed a lot of energy and in case
of feed solution is in food industry, it effected to the quality of the product. However, the
too low feeding temperature is also not acceptable because the difference of temperature
between two sides of membrane is not enough to create gas phase in the membrane
surface. Thus, the best temperature found out. The temperature in membrane surface could
not be measured by the experiment. It was calculated by Equation 2.36, 2.37 for DCMD
configuration and Equation (2.38 – 2.41) for SGMD configuration.
3.3.3 Gas flowrate calibration (in SGMD)
In SGMD, membrane flux closely related to sweeping gas flow. The higher sweeping gas
flow rate bring to increasing permeate flux due to turbulent regime, but when gas flux
reached a certain value, the permeate flux does not increase any more, that mean it is
stable. Thus the excess gas flux is not necessary, so it also lead to consume more energy.
However, the small gas flow rate does not create enough pressure to push the permeate out
of membrane module. The appearance of bubbles illustrated this phenomenon. Taken
together, finding the best value of sweeping gas flux was very importance.
3.3.4 Temperature polarization
The boundary layers at feed and permeate side were formed due to effect of viscosity of
liquid solution and liquid flow rate. This phenomenon was called temperature polarization
that enforced a resistance to heat transfer through boundary layers. Therefore, temperature
49
was difference between membrane surface and bulk side, the membrane flux could be
reduced by this phenomenon as presented in Figure 3.10. Another phenomenon also
affects to membrane flux that was concentration coefficient, but the effect was negligible.
Temperature polarization coefficient was used to evaluate the effect of the presence
boundary layers that was shown in Equation 2.30.
Sweeping Gas/
Cold water
Feed
Tb,f
Tb,p
Tm,p
Tm,f
Qp QfQm
Figure 3.10 Temperature polarization in hollow fiber (Bui et al., 2007)
3.3.5 System verification
3.3.5.1 Pure water test
The testing with distilled water (DI) was conducted to find out the maximum flux of new
membrane. The method was used that was measuring the mass of DI water reduced every
hour. A mount of water was lost in heater was exactly the amount of permeate passed
through membrane. Thus the flux (m3/m2.h) of membrane was calculated as Equation:
𝐽 =
(𝑊1 − 𝑊2)
𝐴𝑡 (3.1)
Where: W2 is mass of water after one hour measured W1, A is membrane surface area, t is
duration, in this case t equal to 1 hour.
From the value of membrane flux, the membrane coefficient was calculated easily via
Equation 2.6 and 2.7, in which membrane resistance was negligible. Therefore membrane
coefficient in DI case could be compared with membrane coefficient in other aqueous
solution to identify fouling phenomena. The operating parameters were showed in the
Figure 3.1.
50
3.3.5.2 Salt rejection
The rejection test was performed by using 1% saline solution. The test was given to ensure
the rejection ability of membrane, in the other words, only volatile compound (water)
passed through hydrophobic membrane and non-volatile compound (salt) retained in the
feed solution. In addition, the transition from liquid phase to gas phase was taken place in
feeding surface membrane and the membrane did not get wet. The percentage of rejection
(R%) was calculated by using following Equation:
𝑅(%) = (
𝐶𝑓 − 𝐶𝑝
𝐶𝑓) × 100 (3.2)
Where Cp is permeate concentration and Cf is feeding concentration.
3.3.6 The concentration of feeding solution
The experiments with feeding solutions (synthetic solution and food liquid from process)
that were conducted to know the system efficiency at different concentration without any
effect from other impurities in solution.
3.3.6.1 Glucose liquid
There were two sources of feeding glucose liquid as synthetic glucose liquid and glucose
liquid from food process. The concentration of feeding liquid glucose was varied in the
range from 35 – 60 oBrix in both configurations.
Synthetic glucose liquid was prepared according to the rules of mixing solid form glucose
in distilled water according to g/L concentration. The formula is shown as following:
𝐶(𝑔/𝐿) =𝑚𝑠𝑜𝑙𝑢𝑡𝑒
𝑉𝑠𝑜𝑙𝑢𝑡𝑖𝑜𝑛 (3.3)
Where msolute is mass of solute (glucose), Vsolution is the volume of distilled water.
Viscosity vs. concentration
Viscosity of a solution closely is dependent on the concentration and size of solid particles.
As the concentration increases, the viscosity is increased because the distance between
neighboring molecules in solution decreases. Therefore, the molecules move much more
slowly that leading to increasing viscosity of solution.
3.3.6.2 Sodium sulfate solution
The experiment was conducted with sodium sulfate concentration varied from 40 – 450
g/L in both configurations.
Solubility vs. temperature
Temperature and pressure directly relate to solubility of a solute. However, MD system
was operated in the atmospheric pressure, therefore, in this case, its influence was
negligible. Most of salt show the proportional relationship between solubility and
51
temperature. Sodium sulfate is an exception called negative solubility (Tun et al., 2005). It
actually becomes more difficult to dissolve at higher temperatures. This trend was
predicted by using Le Chatelier's principle. In the thermodynamic point of view, the
reaction is divided as following:
Exothermic reactor: energy is put out. Heat is considered as product;
Endothermic reactor: energy is taken in. Heat is included as reactant.
Sodium sulfate dissolved in water was defined as an exothermic reactor. Therefore,
according to the equilibrium law, external heat caused the equilibrium to the exothermic
process by moving towards the reactants.
Na2SO4 + H2O Sodium sulfate solution + Heat
3.3.7 Energy consumption
The energy used in experiment almost came from heating and cooling process, a smaller
part used for air compressor and pump. An electricity meter was supplied to measure the
amount of energy consumed by the whole system.
Energy was used in HF DCMD (E1) and Energy was consumed in HF SGMD (E2) were
calculated by following equations respectively:
𝐸1 = 𝐸ℎ𝑒𝑎𝑡 + 𝐸𝑐𝑜𝑜𝑙 + 𝐸𝑝𝑢𝑚𝑝 (3.4)
𝐸2 = 𝐸ℎ𝑒𝑎𝑡 + 𝐸𝑎𝑖𝑟 𝑐𝑜𝑚𝑝𝑟𝑒𝑠𝑠𝑜𝑟 + 𝐸𝑝𝑢𝑚𝑝 (3.5)
From the results, the ratio of energy consumption and membrane flux of both
configurations at different conditions were compared to find out which configuration at
which condition that consume less energy while the flux still remain in a reasonable level.
The energy observed included energy consumed to heating or/and cooling as well as
energy used for pumping.
𝐸𝑛𝑒𝑟𝑔𝑦 𝑟𝑎𝑡𝑖𝑜(𝑘𝑊
𝑘𝑔) =
𝐸𝑛𝑒𝑟𝑔𝑦 𝑐𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 (𝑘𝑊
ℎ)
𝑀𝑒𝑚𝑏𝑟𝑎𝑛𝑒 𝐹𝑙𝑢𝑥 (𝑘𝑔𝑚2 . ℎ) × 𝑀𝑒𝑚𝑏𝑟𝑎𝑛𝑒 𝐴𝑟𝑒𝑎(𝑚2)
(3.6)
3.4 Parameter Analysis
3.4.1 Glucose analysis
Glucose concentration was estimated by DNS method at wavelength dial to 540 nm. The
principal of the method is that spectrophotometer measured the amount of light absorbed.
The concentration of glucose was determined by using a standard curve. The standard
curve was responsible for translate absorbent value into glucose concentration.
52
Figure 3.11 Standard glucose curve
3.4.2 Sodium sulfate analysis
This method was based on the different weight between before and after heated well –
mixed sample at 180 0C.
Table 3.3 Methods of Analysis
Parameter Unit Method Technique Interference/
Remark Reference
TDS mg/L Dry at
180oC
Filter/Oven
/Water bath
Some sulfate,
chloride ion can
be lost at 1800C
(APHA, 2005),
2540 C
3.5 Membrane Cleaning
Membrane after certain time of operation became congestion leading to reducing
membrane flux, this phenomenon called membrane fouling. Therefore, membrane cleaning
was very critical to remove foulents in the membrane surface. Rinsing membrane module
with DI water was implemented as a recoverable method at initial time and it also used to
remove remained chemical after cleaning membrane with chemical. The cleaning process
was carried out in this study is presented in Figure 3.11.
Organic foulent (Glucose) was removed by dilute alkaline solution (NaOH) and diluted
acidic solution, while inorganic foulent (sodium sulfate/ sodium chloride) was washed out
by dilute acidic solution (0.1wt. % Oxalic acid and 0.8wt. % Citric acid) (Guillen-Burrieza
et al., 2014). The duration of cleaning and the concentration of chemical cleaners were
showed as Table 3.4.
y = 6.5333x + 1.8365
R² = 0.9937
0
2
4
6
8
10
0 0.2 0.4 0.6 0.8 1 1.2
Conce
ntr
atio
n (
mg/m
L)
ABS
Standard Glucose Curve
53
Table 3.4 Chemicals Used for Cleaning Membrane
Foulent Chemical/ method Concentration (%) Duration (h)
Glucose
NaOH 1.5 g/L 6
NaClO 2 wt. %
Oxalic acid 0.1 wt. % 6
Citric acid 0.8 wt. %
TDS Oxalic acid 0.1 wt. %
6 Citric acid 0.8 wt. %
The total fouling resistance (Rt) includes membrane resistance (Rm) and fouling resistance
(Rf). Membrane resistance was indicated by using DI for running membrane module.
Fouling resistance consisted of recoverable fouling resistance (Rr) that was evaluated after
washing membrane with DI water and reversible (Rre), irreversible fouling (Rir) were
identified by cleaning membrane with chemical. The total fouling resistance could be
calculated by following Equation:
Rt = Rm + Rf = Rm + Rr + Rre + Rir (3.6)
54
Rinsing membrane with DI water
until reach the neutral pH
New membrane
Run membrane with distilled water
Run membrane with synthetic/ real
Glucose liquid
Rinsing membrane with DI water
Run membrane module with DI
water
Clean membrane with alkaline and
acidic solution
Run membrane module with DI
water
Run membrane with synthetic
Sodium sulfate/ sodium chloride
liquid
Rinsing membrane with DI water
Run membrane module with DI
water
Clean membrane with acidic
solution
Run membrane module with DI
water
Membrane resistance
(Rm)
Membrane resistance
(Rm) + fouling
resistance (Rt)
Recoverable fouling
resistance (Rr)
Reversible and
irreversible fouling
resistance (Rre + Rir)
Rinsing membrane with DI water
until reach the neutral pH
Figure 3.10 Membrane Cleaning Process
55
Chapter 4
Results and Discussion
This chapter includes the experimental results of the performance of membrane distillation
units in both SGMD and DCMD configuration operated with simulated high TDS
wastewater and glucose liquid. The hollow fiber membrane with 0.255 m2 area and 0.45
µm pore size was carried out in the study. The new membrane was verified by salt
rejection test and pure water test to ensure that MD process operate properly. From pure
water flux, experimental membrane distillation coefficient and resistance were evaluated.
Membrane fouling was also considered in this study. The comparison between DCMD and
SGMD in each category of feed solution in term of flux and specific energy consumption
to figure out the suitable configuration was performed for respective application.
4.1 Membrane Distillation System Calibration
System verification tests were conducted in order to consider the capacity as well as the
mechanism of flux transfer on the hydrophobic membrane. Verification test included pure
water test and salt rejection tests.
4.1.1 Rejection test
The hollow fiber membrane distillation rejection test was conducted to check the rejection
capacity with 1% salt solution. The mission of the rejection test is to make sure the
operating mechanism of membrane distillation works well. The general mechanism for all
configurations of membrane distillation is that the hydrophobic membrane allow only
volatile compound pass through membrane as vapor phase and the non-volatile should be
retained in the feed side. Therefore, the amount of salt solute in the feed side remained
constant after operation because salt is non-volatile compound.
The first rejection test was conducted with gas flow rate is 16.6 L/min then the next
experiment with gas flow rate is 25.5 L/min was also considered. The feeding temperature
is 70oC that is the highest temperature in the range of temperature, which will be conducted
in the study. In addition, at higher temperature, the difference vapor pressure between two
side of membrane also increased, so it may exceed the LEP value, easily lead to wetting
membrane pores.
The rejection is calculated as Equation 4.1
𝑅(%) = (
𝐶𝑓 − 𝐶𝑝
𝐶𝑓) × 100 (4.1)
Table 4.1 Experimental Results of the Rejection Tests on HF SGMD
Gas flow
rate
(L/min)
Flux
(kg/m2.h)
Final feed salt
concentration
(ppm)
Final permeate salt
concentration
(ppm)
Rejection (%)
16.6 1.83 84800 2.24 99.99
25.5 2.88 97920 4.23 99.99
As the results showed in Table 4.1, membrane distillation system removed most of salt
component, that mean there was no salt ion penetrate through membrane pores. In the
56
rejection test, the TDS concentration in permeate side can be affected by the sweeping gas,
so this problem was also considered in the rejection test.
Figure 4.1 The rejection result for HF SGMD with gas flow rate 16.6 L/min
Figure 4.2 The rejection result for HG SGMD with gas flow rate 25.5 L/min
As the results presented in Figure 4.1 and Figure 4.2, the flux was reduced after 3 hours of
operation, however, the rejection capacity of membrane still kept constant at 99.99%.
4.1.2 Pure water test
MD coefficient/ resistance were evaluated by pure water test. In addition, the result from
pure water test were compared with flux after conducting the experiment with real solution
for calculating membrane fouling. Deionized water (DI) was considered as feed solution
for pure water test with hollow fiber membrane
0
20
40
60
80
100
1.6
1.8
2
2.2
1 2 3 4
Rej
ecti
on
(%
)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Rejection
0
20
40
60
80
100
2
2.2
2.4
2.6
2.8
3
1 2 3 4
Rej
ecti
on (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Rejection
57
The parameters were varied in pure water test are feed temperature (50 oC, 60 oC, 70 oC)
and the sweeping gas flow rate (16.9, 19.6, 22.5, 25.5, 28.5 L/min). The temperature of
feed liquid together with flow rate of sweeping gas were found as strongly dependent
factors to permeate flux. The membrane flux was exponentially increased with feed
temperature because of the increased vapor pressure associated with an increase in feed
temperature. The gas flow rate was varied in order to find the optimum value. The
increasing gas flow rate resulted in an increase of Reynolds number that was led to
increase membrane flux (Khayet and Matsuura, 2011). In addition, the variation of gas
flow rate was taken place with the careful precaution so that the pressure difference
between two sides of membrane did not exceed LEP value. Operating condition was
selected based on two criteria were membrane flux and energy consumption ratio.
Figure 4.3 Membrane flux variation at different temperature and gas flow rate
Figure 4.3 presents the membrane flux variation due to changing in operating conditions.
Feed temperature strongly positive effect on membrane flux. The flux significantly
increased as increasing feed temperature. The increase of feed temperature from 50 oC to
60 oC enhanced the membrane flux by around 74%. The highest flux fall into the group
with feed temperature is 70 oC with more than 199% higher flux value compared to the
flux at 50 oC. This parameter is also considered as the most important factor that affecting
directly to membrane flux in DCMD configuration. Accompanying the feed temperature,
gas flow rate was detected as the dependent operating parameter that used to control the
flux membrane. Figure 4.3 reveals that at the same feed temperature, the increasing of
sweeping gas flow rate resulted in an observed increasing of membrane flux. The gas flow
rate enhancement promoted the heat transfer coefficient on the permeate side therefore the
effect of temperature polarization was reduced. However, its effect on membrane flux was
not much significant as in the case of feed temperature. For example, at the same feed
temperature is 60 °C feeding, membrane flux increased by only 40% when increasing the
sweeping gas flow rate from 16.9 to 28.5 L / min. (M. Khayet et al., 2012) also concluded
that the feed inlet temperature affects more intensely than higher gas circulation velocity.
The special phenomena was obtained in the test was the insignificantly decreased flux at
high gas flow rate. The flux reduced by 14% when gas flow rate increased from 25.5 to
1
1
1
2
2
2
3
3
3
4
4
4
5
5
5
0
0.5
1
1.5
2
2.5
3
3.5
50
Mem
bra
ne
Flu
x (
kg/m
2.h
)
Temperature (oC)
1 2 3 4 5
60 70
1 - 16.9 L/min
2 - 19.6 L/min
3 - 22.5 L/min
4 - 25.5 L/min
5 - 28.5 L/min
58
28.5 L/min. The slightly membrane flux reduction is observed with further increasing
sweeping gas flow rate by (Basini et al., 1987). The reason can be explained by the
resistance created by the pressure of high gas flow rate in permeate side.
Table 4.2 The Comparison Effect of Feed Inlet Temperature on Pure Water Flux
(PWF) between This Study and Other Authors.
MD
Type Configuration Material
Feed
temperature
(oC)
Feed
velocity
(m/s)
PWF
(kg/m2.h) Reference
FS DCMD PVDF 40 - 70 0.1 3.6 – 16.2 (Jönsson et
al., 1985)
FS DCMD PTFE 40 - 70 - 5.8 – 18.7
(J.
Phattaranawik
et al., 2003)
FS DCMD PVDF 36 - 66 0.145 5.4 - 36 (Yun et al.,
2006)
FS VMD 3MC 30 - 75 - 0.86 – 9.5 (Lawson and
Lloyd, 1996)
FS SGMD TF 200 40 – 70 0.15 7.2 - 36 (M. Khayet et
al., 2000)
HF SGMD PTFE 50 - 70 4.17 1.05- 3.14 This study
Figure 4.4 Energy consumption variations at different temperature and gas flow rate
The consumed energy was calculated by Equation 3.5. The result is presented in Figure
4.4 shows that increasing feed temperature led to spread used energy. The required energy
for the system included energy for circulation pump, air compressor or thermal energy.
However, the utilized energy for heating accounts for by far 90% of total energy
1
1
1
2
2
2
3
3
3
4
4
4
5
5
5
0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1.0
50
Ener
gy c
onsu
mpti
on (
kW
/h)
Temperature (oC)
1 2 3 4 5
60 70
1 - 16.9 L/min
2 - 19.6 L/min
3 - 22.5 L/min
4 - 25.5 L/min
5 - 28.5 L/min
59
requirement (Khayet and Matsuura, 2011). At the same gas flow rate 25.5 L/min, the
energy consumption raised by around 89% when the feed temperature increased from 50 oC to 70 oC. Therefore, it was to note that energy requirement increase drastically with the
enhancement of feed temperature. The energy consumed for supplying gas gained higher at
higher flow rate. Anyhow, the difference in energy consumption for higher gas velocity is
relatively low as observed in Figure 4.4. The energy consumption ratio calculation is
explained in Equation 3.6 revealed that the smaller energy requirement value achieve, the
system is more efficiency. The energy ratio was lowest at 70 oC of feed temperature, so 70 oC was chosen as the best operating temperature. The energy ratio value was relatively
similar at gas flow rate 22.5 and 25.5 L/min as 1.07 and 1.09 respectively. However,
compared to the energy consumption, membrane flux is considered to be more priority
factors. The gas flow rate 25.5 L/min created higher flux by 20% compared with flux at
22.5 L/min flow rate. Thus the 22.5 L/min air flow rate and 70 oC of feed temperature were
selected as operating conditions for the experiments with synthetic and real feed solution.
Figure 4.5 Energy ratio variations at different temperature and gas flow rate
In the desalination process, in term of energy consumption and process, MD also gained a
lot of favor from the scientists compared with conventional processes. Table 4.3 presented
the energy requirement ratio comparison between MD and other conventional methods in
desalination process. The consumed energy of reverse osmosis is quite lower than energy
used in MD process because only mechanical energy is utilized in RO process while both
mechanical and thermal energy are required for MD process. However, in term of process,
RO has big drawback of fouling phenomena due to small size of pores.
1
1 1
2
2 2
3
3
3
4
4
4
5
5
5
0.0
0.5
1.0
1.5
2.0
2.5
3.0
50
Ener
gy c
onsu
mpti
on r
atio
(kW
/ kg)
Temperature (oC)
16.9 19.6 22.5 25.5 28.5
60 70
1 - 16.9 L/min
2 - 19.6 L/min
3 - 22.5 L/min
4 - 25.5 L/min
5 - 28.5 L/min
60
Table 4.3 The Comparison between MD and Conventional Method in Desalination
Process.
Process Energy requirement
ratio (kW/kg) Year Reference
Membrane Distillation 1.85 2008
Adapted from
(Yarlagadda et
al., 2010)
Reverse Osmosis 1.37 2002
Multi Stage Flash Distillation 5.63 1996
Multi Effect Distillation 4 1998
Multi Effect Solar Still 25 2005
HF membrane distillation 1.09 2015 This study
4.1.3 Membrane coefficient and membrane resistance
Membrane resistance (Rw) and membrane coefficient (Bw) was also figured out by pure
water test result. The relationship between membrane resistance and PWF, membrane
resistance and membrane coefficient were established by Equation 4.2 and 4.3
respectively:
𝑅𝑤 =
∆𝑝𝑤
𝐽𝑤 (4.2)
𝐵𝑤 =
1
𝑅𝑤 (4.3)
The bulk feed and permeate temperature were calculated by Equation 4.4, 4.5. In which,
the bulk inlets and outlets were measured by temperature sensor.
𝑇𝑏,𝑓 =
𝑇𝑏,𝑓𝑖𝑛+𝑇𝑏,𝑓𝑜𝑢𝑡
2 (4.4)
𝑇𝑏,𝑝 =
𝑇𝑏,𝑝𝑖𝑛+𝑇𝑏,𝑝𝑜𝑢𝑡
2 (4.5)
The membrane feed and permeate temperature could not be measured in the experiment,
instead, they were calculated based on the theoretical equations from (Martı́nez-Dı́ez and
Vázquez-González, 1999) for membrane feed and equations from (Khayet and Matsuura,
2011) for membrane permeate temperature of sweeping gas. Membrane feed temperature
was obtained from following procedure:
Assume initial value of membrane feed temperature;
Calculate the vaporization heat Hv by Equation 4.6
∆𝐻𝑣,𝑤 = 1.7535𝑇 + 2024.3 (0K) (4.6)
While T= 𝑇𝑚,𝑓+𝑇𝑚,𝑝
2
Calculate Tm,f and Tm,p by Equation 2.36, 2.37 and the result from Equation 4.6
The procedure was repeated until the calculated membrane temperature and assumed ones
difference had error of less than 0.1%.
61
The membrane permeate temperature was calculated by Equation 2.39 – 2.41.
The Table 4.4 expressed the result of calculated feed, permeate temperature and
temperature polarization coefficient (TPC). Detail of the calculation procedure is presented
in Appendix C4
Table 4.4 The Membrane Surface Temperature and TPC
Feed (oC) Permeate (oC) TPC
hf W/(m2.k) Tb,f Tm,f hp W/(m2.k) Tb,p Tm,p
420.4 69.5 67.7 163.6 43.5 60
0.3 418.5 59 56.7 167.17 43 52.4
413.5 49 47.5 167.52 38 43.7
The permeate heat transfer coefficient is found much smaller than in the feed side. In
addition, the temperature in the bulk feed and membrane feed is relatively similar, TPC in
feed side approaches to unity point. Therefore, the TP is mainly localized in the permeate
side of SGMD process. In other words, the permeate heat transfer resistant control the TPC
of process. According to the research of (Khayet et al., 2002) with plate and frame SGMD
system, TPC is obtained less than 0.44. They also indicated that this value is chiefly
contributed by the permeate TPC.
Table 4.5 Membrane Resistance and Membrane Coefficient Calculation Value in
SGMD Configuration
Membrane Feed
Temperature (oC)
Membrane Coefficient
(×𝟏𝟎−𝟖 s/m)
Membrane resistance
(×𝟏𝟎𝟔 m/s)
Hollow fiber
PTFE
0.45 µm
49 2.8 35
59 2.9 34
69.5 4.56 22
The membrane resistance and coefficient are calculated in Table 4.5. Membrane resistance
has a negative effect on membrane flux. The enhancement of membrane resistance
accompanied with reducing feed temperature. (Srisurichan et al., 2006) also asserted that
the increasing feed temperature contributed in decreasing of membrane resistance.
4.2 TDS Removal Test
TDS removal test were conducted on hollow fiber membrane with pore size 0.45 µm,
packing density is 333.8 m2/m3. Two configurations that were considered in the experiment
are SGMD and DCMD. It is important to ensure that in operation of the system, the
pressure in the feed side of membrane (caused by feeding flow) and the pressure in
permeate side (caused by the flow of liquid or gas flow) were relatively equal. If the
pressure difference between two sides of membrane exceeds the LEP value, the wetting
pore phenomena will occur. Moreover, the main mechanism of membrane distillation was
the vapor pressure difference between two sides of membrane, rather than different
pressure caused by such flow rates. Membrane distillation process operated at atmospheric
pressure. The membrane was tested with divalent salt solution to examine the performance
of membranes working with only salt. The desired concentration of divalent salt solution
was prepared by mixing Na2SO4 with DI water.
62
4.2.1 TDS removal on hollow fiber sweeping gas membrane distillation HF-SGMD
The experiments on SGMD system were conducted under the optimum condition that was
selected from results obtained from system calibration test (Section 4.1). The operating
conditions were as follows; feeding temperature 70 oC, permeate gas (ambient air) at
ambient temperature (25 -300C), sweeping gas flow rate of 25.5 L/min and feed flow rate
of 2.4 L/min.
4.2.1.1 Testing the capacity of membrane with high concentration salt solution in HF
SGMD
In this section, experiments were conducted using continuously feeding method to
determine the maximum obtainable rejection and flux. Divalent salt solution was added
into the feed tank after every 8 hours with the concentration of the next batch 10 % lower
than the end concentration of previous batch. In addition, the volume of the feed solution
was maintained at 10 liters while starting a new batch. Sodium sulfate was very prone to
crystallize at high concentrations due to its negative solubility at higher temperature.
Accumulated crystal salt could negatively affect the hydrophobic characteristic of
membrane therefore the membrane was rinsed with distilled water after each batch to
ensure the salt crystals could not form when the system was turned off. In this study, the
calculation of fouling resistance would be limited because of rinsing membrane with DI
water. However, the membrane flux achieved the initial value after the membrane was
cleaned follow the washing procedure with dilute acidic solution as presented in Section
3.5. The procedure of rinsing was also confirmed by researchers using DCMD with
sparingly soluble salt (CaSO4), the effective control of CaSO4 scaling was controlled by
simple regular flushing membrane with water (Nghiem and Cath, 2011).
Figure 4.6 Flux and concentration with high concentration Na2SO4 solution in HF
SGMD
As presented in Figure 4.6, the permeate flux did not significantly change when the
concentration of salt was increased. The flux decreased slightly from 3.07 kg/m2.h to 2.23
kg/m2.h when the salt concentration increased from 40 g/L to around 450 g/L (Table 4.6)
as the salt solution had a low viscosity therefore the effect boundary layer and its related
resistances were minimal. In addition, concentration and temperature polarization
phenomena suggests that the temperature was lower and salt concentration was higher on
0
200
400
600
800
0
1
2
3
0 5 10 15 20 25 30
Conce
ntr
atio
n (
g/L
)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Concentration
63
the membrane surface than observed in the bulk side of membrane (Ali et al., 2013). On
the other hand, solubility of the salt has a close relationship with temperature. Depending
on the type of salt (positive solubility or negative solubility), the solubility is proportional
or inversely proportional to the temperature (Boundless, 2015).
Table 4.6 Summary of Na2SO4 Solution at Optimum Conditions in SGMD
Batch
Theoretical Na2SO4 concentration
(g/L)
Analysis Na2SO4 concentration
(g/L)
Initial Final Initial Final
1 40 103 38.8 99.5
2 93 246 73.9 201.1
3 221 494 196.2 392.7
4 445 749 359.8 457.8
Sodium sulfate has negative solubility, implying higher solubility at lower temperature as
presented in Figure 4.7. In this case, temperature polarization had a favorable effect, in
other words, the solubility of Na2SO4 salt solution was higher at the membrane surface
than in the bulk feed (Chernyshov et al., 2003). Therefore, the salt crystal did not formed in
the membrane surface before it appeared in the bulk feed. The process was conducted in
nearly 30 hours. The test reached around 450 g/L then some salt crystal started to form and
accumulated to the membrane surface. The reason for this phenomenon was that at 70 0C,
only about 40 grams of Na2SO4 salt dissolved in 100 g of water, or in other words, the
solution was saturated as seen in Figure 4.7.
Figure 4.7 Solubility of sodium sulfate vs. temperature (adapted from (Linke., 1958))
As seen in the tail of the Figure 4.6, the concentration of feed solution reached super
saturation point which was predicted in a drastic membrane flux decline because of the
rapid forming of salt crystal on the surface of membrane (Tun et al., 2005). The result from
Figure 4.8 that was also mentioned in Section 3.3.7 shows that as low concentration of
divalent salt solution, the required energy consumption/ membrane flux ratio for heating
was stable at ~1.07 kW/kg, because the viscosity of solution was low or in other words, the
energy consumption for latent heat was remained at higher concentration of salt solution.
0
10
20
30
40
50
20 25 30 40 50 60 70 80 90 100
Solu
bil
ity/
gra
m p
er 1
00 g
H2O
Temperature (oC)
Sodium Sulfate Solubility
64
However, the required energy ratio tends to increasing when solution reached saturation
point.
At the saturated point of Na2SO4 solution, the permeate flux reduced. However, the
required energy kept constant therefore, the consumed energy ratio was significantly
increased to 1.97 kW/ Kg. Although the consumed energy was supposed to reduce with the
reduced quantity of feed solution with time. However as the solution had a low viscosity,
thus the ability of the solution to retain heat was low. In addition to low viscosity, energy
consumption of the pump kept constant as pump needs more energy to operate with high
viscosity solutions.
Figure 4.8 Energy consumption ratio of testing the capacity of membrane with high
concentration Na2SO4 solution in HF SGMD
4.2.1.2 Testing the membrane with high concentration salt solution in real operation
The test was conducted to simulate real wastewater. As the volume of the feed tank was
very small thus the volume of solution need to be handled in batch operations. The initial
concentration of feed solution was 40 g/L, operation time was 8 hours per batch then the
feed tank also was filled up by 40 g/L of Na2SO4 solution. The experiment was considered
in a long process so as to increase the concentration of feed solution without any gaps from
40 g/L to approximately 500 g/L. The overall flux and sodium sulfate concentration with
time is presented in Figure 4.9.
0
200
400
600
800
0.0
0.5
1.0
1.5
0 5 10 15 20 25 30
Conce
ntr
atio
n (
mg/L
)
Ener
gy c
onsu
mpti
on r
atio
(kW
/Kg)
Time (h)
Energy Concentration
65
Figure 4.9 Flux and concentration with high concentration Na2SO4 solution in HF
SGMD during simulated real operation
The permeate flux of the process was relatively stable around 2.51 kg/m2.h while the
concentration of the feed solution increased from 40 g/L to ~ 450 g/L (Table 4.7). This
was due to the TDS solution had a low viscosity therefore it had less effect to forming
boundary layer resistance. In addition, the temperature polarization had positive effect on
the solubility of Na2SO4 salt solution. However, the system could only operate with
maximum concentration of Na2SO4 solution around 450 g/L because at this point some salt
crystal started to accumulate in membrane surface even at high temperature (70oC). Khayet
et al. (2003) also observed similar results and concluded that the flux in SGMD decreased
slightly with an increase in sodium chloride concentration for desalination.
Table 4.7 Theoretical and Measured Concentration of Na2SO4 Solution with HF
SGMD Simulating Real Operation for Phenol Industry Wastewater
Batch Theoretical Na2SO4 concentration
(g/L)
Analysis Na2SO4 concentration
(g/L)
1 40 41.8
6 280 245.9
9 486 452.7
Even with batch feeding operation, the energy consumption for the system remained
unchanged at ~1.07 kW/kg while increasing concentrations. But it could be noted that an
increase in energy consumption was observed when the concentration was close to
saturation point. Energy ratio did not change compared to previous experiment, it slightly
fluctuated between 1.02 – 1.6 kW/ kg. According to (Criscuoli et al., 2008), the ratio of
energy consumption and permeate flux in vacuum membrane distillation configuration was
1.1 Kw/kg. In aggregate, HF SGMD could operate with extremely high divalent salt
solution at reasonable flux. In case of dealing with negative solubility salt (Na2SO4),
temperature polarization was observed to be favorable for enhancing membrane flux until
feed solution reached the saturation point.
0
100
200
300
400
500
0.0
0.5
1.0
1.5
2.0
2.5
3.0
3.5
0 5 10 15 20 25 30 35 40 45 50 55 60 65
Conce
ntr
atio
n (
g/L
)
Flu
x (
kg/m
2.h
)
Time (h)
Permeate flux Concentration
66
Figure 4.10 Energy consumption ratio of membrane with high concentration salt
solution in real operation in HF SGMD
The temperature polarization coefficient of SGMD in operating with highest concentration
of divalent salt solution was presented in Table 4.8. TPC value was remained at the same
value that calculated from experiment in SGMD with pure water. Therefore, the salt
concentration had negligible effect on TPC value. This phenomenon also reinforced
opinion presented in Section 4.1.3 that in SGMD, temperature polarization is stationed in
permeate side. Therefore, the more significantly effect on TPC will be occurred in case of
changing the sweeping gas flow rate.
Table 4.8 Temperature Polarization Coefficient in SGMD Configuration with High
Concentration of Salt Solution
Feed (oC) Permeate (oC) TPC
Bulk Membrane Bulk Membrane
68 58 43 50.5 0.3
4.2.2 TDS removal on hollow fiber direct contact membrane distillation HF-DCMD
The experiments on DCMD system was conducted under the same condition that was
chosen in SGMD system as feeding temperature of 70 oC, feeding flow rate of 2.4 L/min.
The permeate temperature was kept at 10 oC, however, the real temperature was operated
in the system fluctuated in the range 11 - 17 0C due to heat exchange and conduction
through the membrane surface and also loss of heat with environment. The permeate flow
rate in DCMD play an important role for adjusting the membrane flux. The enhancement
permeate flow rate resulted in higher membrane flux due to minimize the negative effects
of temperature polarization. However, the flow rate in permeate had to relatively equal
0
100
200
300
400
500
0.0
0.5
1.0
1.5
2.0
0 5 10 15 20 25 30 35 40 45 50 55 60 65
Conce
ntr
atio
n (
mg/L
)
Ener
gy R
atio
Consu
mpti
on (
kW
/Kg)
Time (h)
Energy Consumption Concentration
67
with flow rate in feed side to ensure that the pressure difference did not appear that
affected the selectivity nature of hydrophobic membrane. The LEP value of the membrane
could be calculated by Equation 2.1 as following:
𝐿𝐸𝑃 =−2𝐵𝛾𝑐𝑜𝑠𝜃
𝑟𝑚𝑎𝑥
Where,
Cylindrical pore has B = 1 (Rácz et al., 2014)
The surface tension is 72 mNm-1 for pure water, while, inorganic salt solution has greater
surface tension than pure water. Therefore, the surface tension of sodium sulfate solution is
assumed as 72 mNm-1 as safety factor.
so γ > 72 mN/m
Contact angle θ = 112o (PTFE membrane property)
Maximum pore size rmax = 0.45 µm
Therefore:
𝐿𝐸𝑃 =−2 × 1 × 72 × 10−3 cos(112)
0.45 × 10−6= 119.9 (𝐾𝑃𝑎)
The flow rate in feed side was selected (~2.4 L/min) to make the system comparable to HF
SGMD in Section 4.2. Therefore, the cooling water flow rate in permeate side was also
maintained as close as possible to this value. However, energy consumption factor was
also taken into consideration. The pump used in the system could be adjusted with several
testing condition as seen in Figure 4.11 below.
Figure 4.11 Performance of permeate pump in HF DCMD system at different flow
rate
As presented in Figure 4.11 there was nearly no increase in flux with higher permeate flow
rate after 1.32 L/min. According to Ohta et al. (1990), the membrane flux increased with
the cooling water flow rate. This was also demonstrated in the study of (Close and
Sørensen, 2010), cool water flow rate increased leading to increased membrane flux,
however, the increase in membrane flux gradually decreased with increasing cooling water
2.61 2.61
2.88 2.88 2.88 2.88
2.88
2.5
2.6
2.7
2.8
2.5
2.6
2.7
2.8
2.9
3
0.74 1.06 1.32 1.69 2.1 2.51 2.68
Ener
gy c
onsu
mpti
on (
kW
/h)
Mem
bra
ne
Flu
x (
Kg/m
2.h
)
Flow Rate (L/min)
Energy Flux
68
flow rate. The flow rate 1.32 L/min was chosen as it created higher flux with relative lower
energy consumption amongst all the tested permeate flow rates.
The membrane coefficient and membrane resistance in DCMD configuration were also
determined by Equation 4.2 – 4.6
Table 4.9 Membrane Resistance and Membrane Coefficient Calculation Value in
DCMD Configuration
Feed (oC) Permeate (oC) TPC
Membrane
coefficient
Membrane
resistance Bulk Membrane Bulk Membrane
63 53.8 17 19.9 0.74 6.4 × 10-8 15.6 × 106
The TPC value of membrane in DCMD configuration is approximately close to unity. High
TPC indicated DCMD value is well design. However, according to (Khayet and Matsuura,
2011), the mass transfer is limited because of low membrane permeability when TPC
value in DCMD module is greater than 0.6.
Table 4.10 Membrane Coefficient Comparison between the Study with some
Researchers in DCMD Configuration
Membrane
Type
Pore size
(µm)
Membrane Coefficient
10-8 (s/m) Ref.
PTFE 0.45 21.5 (García-Payo et al., 2000)
PVDF 0.22 3.8 (Imdakm and Matsuura, 2004)
PVDF 0.45 4.8 (Zolotarev et al., 1994)
HVHP 0.45 6.614 (Martı́nez et al., 2003)
PTFE 0.45 6.4 This Study
Table 4.10 revealed that the membrane coefficient of the study is quite high compared to
other researches.
After selection of the appropriate flow rates, the experiment was conducted to determine
the performance of hollow fiber DCMD by the approach of continuously feeding. The feed
tank was filled by divalent salt solution after every 8 hours with the calculated
concentration of the next batch was lower than the end calculated concentration of
previous batch 10%. The volume of feed solution was always kept at 10 liters in the
beginning of new batch. The membrane also was rinsed by DI water after each batch to
prevent affecting from salt crystal to membrane surface.
69
Figure 4.12 Performance of HF DCMD in the operation with high concentration
Na2SO4 solution
The HF DCMD system had the ability to concentrate salt solution from 40 g/L to around
450 g/L similar to SGMD tested in the previous section. Figure 4.13 reflects the similar
between theoretical TDS value and measured TDS value. As seen in Figure 4.12 following
at the salt solution concentration ~ 300 g/L the flux slightly decreased from 3.07 to 2.51
Kg/m2.h, while with the salt concentration was higher than 300 g/L and came closer with
saturation point, the flux was reduced more significantly from 2.51 to 1.12 Kg/m2.h. While
concentration of Na2SO4 solution was ~450 g/L, the concentration in membrane surface
reached the saturation point therefore salt crystals were observed on the membrane surface.
Table 4.8 presents the data for the salt concentration for all the six batch operations. The
greatly difference in concentration between bulk and membrane surface was the underlying
cause leading to a significant reduction in membrane flux. In the entire of investigated
range, the results of experiment showed a sharply decline of membrane flux of 63.5%.
Cath et al. (2004) worked with NaCl in DCMD system, at feed and permeate temperature
40 and 20 oC, respectively. The researchers concluded that an increase in concentration of
salt solution had a minor effect on flux with only 9% flux decline over the experimental
period. However their research was conducted at 0.6 to 73 g/L of NaCl which was much
lower than this study. In another research by Yun et al. (2006) investigated DCMD system
with high concentration of NaCl solution, the feed and permeate temperature respectively
are 79 and 20.5 oC, their result shows that along with the increase in concentration,
membrane flux changes as following three stages, firstly, the flux decrease slightly with
time after that the membrane flux drop dramatically until the salt concentration reaches
saturation point, the flux keeps stabilizing. In the whole process, membrane flux decreases
by more than 90%.
0
100
200
300
400
500
600
700
800
0.0
0.5
1.0
1.5
2.0
2.5
3.0
3.5
0 5 10 15 20 25 30 35 40 45 50 55
Con
centr
atio
n (
mg/L
)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Concentration
70
Figure 4.13 Measured concentration of Na2SO4 solution for HF DCMD in the
operation with high concentration Na2SO4 Solution
According to (Nghiem et al., 2011), the experiment in DCMD system, the sudden drop of
permeate flux was observed when the feeding solution was saturated calcium sulfate
solution. Their research also indicated that the contact angle value of membrane was
reduced after the experiments. This phenomenon revealed the decline of hydrophobic
nature of membrane.
Figure 4.14 Energy ratio with increasing salt concentration in HF DCMD
The heater and cooler equipment represented as main energy requirement sources of the
DCMD process. As presented in Figure 4.14, required energy for DCMD was
unchangeable at low concentration of sodium sulfate and average around ~3.6 Kw/kg.
Criscuoli et al. (2008), presented while using DCMD configuration the lowest required
energy required for their system was at 3.55 kW/kg, where feed temperature varied
between 40 – 60 oC while permeate was between 13 – 14 oC. However, the required energy
tended to greatly increase and rose up to 60 % of energy demand compared to working
with the initial concentration. The possible reason for this of extremely change energy ratio
0
200
400
600
800
0 10 20 30 40 50
Sal
t co
nce
ntr
atio
n(
g/L
)
Time (h)
Theoretical TDS
Measured TDS
0
100
200
300
400
500
600
700
800
0
2
4
6
8
10
0 5 10 15 20 25 30 35 40 45 50 55
Conce
ntr
atio
n (
mg/L
)
Ener
gy R
atio
(K
w/k
g)
Time (h)
Energy Concentration
71
was the flux reduction at high salt concentration, not due to the decreasing requirement of
energy for heating or cooling aqueous solution.
In DCMD configuration, the boundary layers are recognized as the factor that limits the
heat transfer or in other words, cause of constraining membrane efficiency.
Table 4.11 Temperature Polarization Coefficient in the Test of Salt Solution with
DCMD
Feed (oC) Permeate (oC) TPC
hf W/(m2.k) Tb,f Tm,f hp W/(m2.k) Tb,p Tm,p
100.2 66 41.5 1342.1 18 19.8 0.45
Generally, the TPC value is in the range of 0.4 – 0.7 is considered as tolerable DCMD
module (Khayet and Matsuura, 2011), therefore, TPC of DCMD with high salt solution
concentration ~ 0.45 is satisfactory. The result of low feed heat transfer value in Table
4.12 indicated that the dominant reason of heat transfer resistance is comprised from feed
side. However, in case of operating with high concentration salt solution, the lower TPC
value of membrane also asserted that the membrane temperature difference between two
sides of membrane was reduced because of boundary layers effect. This is a major reason
of decreasing membrane flux at high concentration of salt.
4.2.3 Fouling analysis in MD with high concentration of salt solution
In the experiment with synthetic solution, MD process is generally affected by membrane
resistance and boundary layers resistance. The membrane resistance is constant at 21.9 ×
106 m/s that was evaluated in the pure water test, while boundary resistance closely
depended on the concentration of feed solution.
Table 4.12 Membrane Resistance and Boundary Layers Resistance.
Salt
concentration
(g/L)
Flux
(kg/m2.h)
Membrane
Resistance
(m/s)
Feed boundary
layer resistance
(m/s)
Permeate
boundary layer
resistance (m/s)
0 3.14
21.9 × 106
26.29 × 105 56.68 × 105
50 3.07 26.89 × 105 57.27 × 105
150 2.51 32.89 × 105 62.56 × 105
250 2.51 33.83 × 105 62.29 × 105
350 2.23 37.02 × 105 65.53 × 105
450 1.95 42.33 × 105 69.07 × 105
Table 4.13 presented that the higher concentration of salt facilitated to the growth of
boundary layers resistance. As the salt concentration increased from 50 to 450 g/L, the
contribution of boundary layers resistance increased from 27.5% to 33.7 %, in addition, the
flux reduced 36.5%. Further analysis revealed that membrane resistance accounted as the
highest resistance that negative affect to membrane flux. The slightly increase of feed
boundary layer is quite difficult to observed in Figure 4.15 prove that the enhancement of
salt concentration in feed side had mildly impact to boundary resistance, in other words,
the salt concentration had insignificantly effect on membrane flux. The effect of salt
concentration on the membrane flux even was not more significantly than effect caused by
72
the low thermal conductivity of the gas in the permeate, which reflected in higher rates of
permeate boundary layer resistance than its in feed side.
Figure 4.15 Ratio of membrane resistance and boundary layers resistance in MD with
high salt concentration solution
Fouling analysis in MD with high concentration of salt solution was conducted to consider
the percentage of irreversible fouling. The fouling analysis is similar between DCMD and
SGMD configuration because the salt concentration and contact time did not affect very
seriously on the fouling of negative solubility salt. The inorganic fouling was generated
due to the formation of sodium sulfate crystallization. The crystal nucleation appeared at
supersaturated condition. The different resistances contributions are grouped in Table 4.14
and Figure 4.16
Table 4.13 Different Type of Resistance in MD with High Sodium Sulfate Solution
Resistance Value (106 m/s) Percentage
Membrane 22.55 44.7
Boundary layers 11.14 22.1
Inorganic
Fouling
Recoverable 12.85 25.5
Reversible 3.89 7.7
Irreversible 0 0
Total 50.43 100
The highest resistance accounted for 45% that was localized in membrane resistance while
the inorganic fouling did not play an importance role in resistance even at saturated point
of salt solution. Negative solubility of sodium sulfate salt is rational reasons to explain this
phenomenon. Negative solubility salt is more soluble at low temperature that means the
crystal was flimsily formed in membrane surface. However, homogenous crystallization
nucleated in the bulk feed was also presented as the further thermal resistance led to flux
reduction that was showed in Section 4.2.1 and 4.2.2. Additional, raising membrane with
DI water in the end of each batch is the reason for superficial influence of inorganic
fouling. The membrane was relatively recovered completely by running with DI water for
0
0.5
1
1.5
2
2.5
3
3.5
40%
50%
60%
70%
80%
90%
100%
0 50 150 250 350 450Salt solution (g/L)
Flu
x (
kg/m
2.h
)
Membrane resistance Feed Boundary resistance
Permeate Boundary resistance Flux
73
around 30 min with recoverable fouling accounted for 25.5%. Some researchers conducted
the recoverability of MD is presented in Table 4.15.
Table 4.14 Recoverability of MD in Some Research
Membrane
material
Feed solution Cleaning method Recovery
(%)
Reference
PTFE 7g/L NaCl Water wash 98.48
Adapted from
(Warsinger et al.,
2015)
PP Ground water
(CaCO3) 3 wt.% HCl 98.75 (Gryta, 2010)
PP Ground water
(CaCO3, CaSO4) 2.5 wt.% HCl ≈100 (Gryta, 2008)
PTFE
450 g/L Na2SO4 Water wash 92.3
This study 0.1%w.t Oxalic,
0.8%w.t Citric ≈100
Figure 4.16 Different type of resistances in MD with high salt concentration
4.3 Glucose Liquid Concentration Test
Bench scale glucose liquid concentration test was carried out in both DCMD and SGMD
configuration. Synthetic glucose liquid and filtered real glucose liquid were used as feed
solution to considering the performance of those configurations.
Glucose liquid is considered as a thermally sensitive solution. Therefore, liquid glucose
was operated at low temperature. The excess heat on glucose liquid in evaporation process
can cause the risk of color formation. In the research of (Bui et al., 2007) to concentrating
glucose liquid in DCMD, low feed temperature of 40 oC is selected to conduct the
experiment. However, the quality of glucose liquid will not be affected at 50, 60 oC as
mentioned in the recommended storage temperature of glucose liquid (Hull, 2010). In
addition, energy consumption factor was also taken into consideration, thus the
Membrane
45%
Boundary
layers
22%
Recoverable
25 %
Reversible
8 %
Irreversible
0 %
Inorganic
Fouling
33%
74
temperature was used in the experiment of the study is 50 oC for both DCMD and SGMD
configuration.
4.3.1 Glucose liquid concentration on hollow fiber direct contact membrane
distillation (HF DCMD)
The experiment of concentrating glucose liquid on DCMD was operated under the
condition that was used in TDS removal test except feed temperature. The experimental
condition was 50 oC for feed temperature, 10 oC for permeate temperature, 1.32 L/min for
cool water flow rate, and 2.4 L/min for hot glucose liquid flow rate.
4.3.1.1 Synthetic glucose liquid concentration on HF DCMD
The synthetic glucose liquid with desired concentration was performed by dissolving raw
glucose with deionized water. The initial feed solution was synthetic glucose liquid with 10
L of 10% w.t glucose concentration. The feed tank was always maintained at 10 L by
adding glucose liquid. The starting concentration of next batch was 10% lower than the
final concentration of previous batch. Each batch was conducted for 12 hours. The
membrane was only cleaned in the end of experiment. Boundary layers resistance grew up
in membrane surface thus fouling analysis was included in this section.
The HF DCMD system was able to concentrating glucose liquid up to 74% w.t. The
experiment consumed nearly 450 hours to concentrating glucose liquid from 10% to
approximately 75%.The similarity of measured glucose concentration by DNS method and
theoretical value is expressed in Figure 4.18. The preliminary statistics from Figure 4.17
shows that DCMD flux of 30% glucose solution was able to reach around 62.7% of pure
water flux. (Schofield et al., 1990) also stated that the DMCD system operated with 30% of
sucrose solution achieves 60 - 70% of the membrane flux is dealt with DI water. The
permeate flux reached 10% of initial flux was the signal for finishing the experiment.
(Curcio et al., 2000) described that the apple juice could be concentrated up to 64 oBrix by
using HF DCMD (membrane area is 0.1 m2, pore size is 0.45 µm). The flux achieved from
their experiment was 1–1.5 kg/m2.h. A further assertion for the ability to concentrate food
product (glucose liquid) of MD was published by (Bui et al., 2007). They reported that
glucose liquid concentration is able to reach 60% by lab scale HF DCMD system. The
functionality of membrane flux and concentration versus time is presented in Figure 4.17.
In the initial, the flux was fluctuated between 1.61 and 1.21 kg/m2.h after that, the gradual
flux reduction was observed as increasing glucose concentration. The study of MD
performance with sugar solution and orange juice also endorsed that membrane flux mildly
diminished with an increase feed concentration (Rodrigues and Fernandes, 2012). Initially,
membrane flux was not stable; it fluctuated from 1.61 to 1.21 kg/m2.h. After that the flux
gently reduced from 1.61 kg/m2.h at glucose concentration 10% w.t to 0.2 kg/m2.h at
highest glucose concentration (74% w.t). Thus, it was supposed that glucose concentration
imposed a negative effect on membrane flux performance. As shown in Figure 4.17, trans-
membrane flux steadily decreased with the feed glucose liquid concentration growth. There
was no significant change in membrane flux as increasing concentration. Adapted
information from the research of (Lawson and Lloyd, 1997), there was three possible
reasons to explain the flux reduction. Firstly, water activity (aw) decreases when the feed
concentration (x) increases thus reducing vapor pressure in feed solution. This theory is
proven by Equation 4.7.
75
𝑝𝑤(𝑥, 𝑇) = 𝑝𝑤0 × 𝑎𝑤 (4.7)
The second and third reason are mass transfer resistance and heat transfer resistance due to
the development of boundary layer as higher feed concentration. In addition, the flux
diminution as a consequence of increasing the glucose concentration was also attributed by
the upgrade of glucose liquid viscosity. Viscosity of sucrose was also indicated as the
dominant factor leading to flux reduction in the research of (Schofield et al., 1990) on
DCMD configuration.
Figure 4.17 Permeate flux vs. feed synthetic glucose liquid in HF DCMD system
However, an observation in (Rodrigues and Fernandes, 2012) research of comparison
between MD and RO process indicated that the higher flux is attained in MD than flux in
RO process for concentrating orange juice. Specifically, the initial flux of RO is
significantly higher than MD flux. Nonetheless, the RO flux is immediately dropped at
greater orange juice concentration while the flux in MD system just slightly reduce.
(Petrotos and Lazarides, 2001)’s research has shown disadvantage of the direct osmotic
concentration (DOC) system when conducting experiments of concentrating grape juice
solution from 16 to 60 oBrix. The obtained DOC flux is quite high at 2.5 L/m2.h with
cellulose acetate flat sheet membrane. However, the presence of some salt that is diffused
through the membrane is detected in the concentrated grape juice. Saturated brine solution
is utilized as osmotic agent considered as insecure factor in food production process,
whereas, (Bui et al., 2007) implied that the appearance of MD became a fascinated
concentration technology when the quality of the product is priority.
Table 4.15 Comparison between Evaporation and Membrane Techniques
Process
Achieved
concentration oBrix
Product
quality
Flux or
evaporation
rate (L/m2.h)
Energy
consumption Reference
Evaporation 80 Very poor 200 – 300 l/h Very high
Adapted
from (Jiao et
al., 2004)
RO 25 - 30
Very
good 5 - 10 Low
Direct
osmosis 50 Good 1 - 5 Low
MD 60 - 70 Good 1 - 10 Low
DCMD 73% - 0.2 kg/m2.h 27.3 kW/kg This study
0
20
40
60
80
0.0
0.4
0.8
1.2
1.6
2.0
0 50 100 150 200 250 300 350 400 450
Conce
ntr
atio
n (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Concentration
76
Figure 4.18 Measured synthetic glucose liquid concentration by DNS method in
DCMD configuration
The transport resistance enhancement was generated due to the expansion of temperature
and concentration polarization into the membrane surface layers. The transport limitation
alluded to the influence of temperature polarization because concentration polarization was
not play a powerful role in MD. Table 4.18 presents the heat transfer resistance and TPC
value at the highest glucose concentration obtained from the system.
Table 4.16 Temperature Polarization in HF DCMD with Pure Water and Synthetic
Glucose Liquid
Feed solution
Feed (oC) Permeate (oC)
TPC hf
W/(m2.k) Tb,f Tm,f
hp
W/(m2.k) Tb,p Tm,p
Pure water 420.4 48 41.0 1342.1 13 15.1 0.74
Glucose liquid 276.5 48 38.4 1328.1 13 15.0 0.67
The result from Table 4.18 announced that the reduction of TPC value come from the
decrease of heat transfer coefficient in feed solution. The higher concentration of glucose
liquid led to increasing viscosity, in other words reducing Reynolds number that is the
fundamental reason of feed transfer coefficient decline. (Gryta, 2005) reported that
membrane resistance control the MD process when operating with sugar solution
concentration lower than 40% w.t. However, at higher sugar concentration, the
enhancement of feed boundary layer resulted in mass transfer limitation proved to be
dominant. In this case, high membrane coefficient does not play an important role. Figure
4.19 apparently indicates the relationship between membrane resistance and boundary
layers resistance at higher glucose concentration.
0
20
40
60
80
0 5 10 15 20 25 30 35 40
Glu
cose
conce
ntr
atio
n(%
)
Batch (8 h)
Theoretical Glucose Concentration
Measured Glucose Concentration
77
Figure 4.19 Ratio of membrane resistance and boundary layers resistance in MD with
synthetic glucose liquid
Energy consumption is an appreciative point of the MD system. As presented in Figure
4.20, the specific energy consumption gradually grow up with the increasing of synthetic
glucose concentration. Overall view, the energy consumption ratio sharply developed with
82% higher at 74% glucose concentration than its at 10%. The flux reduction at greater
glucose concentration resulted in escalation of energy consumption ratio. Actually, the
energy consumption is quite stable at around 1.85 kW/h because the energy used to re-heat
feed solution kept constant. (Alklaibi, 2008) supposed that economically acceptable MD
system is necessary to obtain the specific energy consumption lower than 50 kW/m3. This
study is profitable because at highest glucose concentration, the energy consumption ratio
reached only around 27.3 kW/kg. Attractive point in concentrating glucose liquid by MD
system is low feed temperature. This is not only advantageous in ensuring the product
quality but also positively affects the energy consumption.
Figure 4.20 Specific energy consumption in HF DCMD system with synthetic glucose
liquid
4.3.1.2 Real glucose liquid concentration on HF DCMD
Filtered glucose liquid 32% concentration obtained from Ajinomoto Company was stored
in cooling room at 4 oC to minimize the fermentation process. The experimental procedure
was similar with the experiment conducted in synthetic glucose liquid with feed solution
0
0.5
1
1.5
2
70%
80%
90%
100%
10 24 34 48 58 69 75Glucose concentration (%)
Flu
x (
kg/m
2.h
)
Membrane resistance Boundary layers resistance Flux
0
20
40
60
80
0
5
10
15
20
25
30
0 100 200 300 400 500
Conce
ntr
atio
n (
%)
Ener
gy r
atio
(kW
/kg)
Time (h)
Flux Energy consumption ratio
78
was filtered glucose liquid 32%. Fouling analysis also was figured out after cleaning
membrane with DI water for recovery resistance, alkaline solution for organic fouling and
acidic solution for inorganic fouling removal.
Figure 4.21 Permeate flux vs. feed real glucose liquid in HF DCMD system
The DCMD configuration was able to concentrate real glucose liquid from 32% to ~60%.
The flux of DCMD with real glucose liquid (0.81 kg/m2.h at 32%) was lower than the
DCMD flux with synthetic ones (0.94 kg/m2.h at 32%). General overview, the
enhancement of real glucose concentration also resulted in reducing membrane flux
significantly. An ~83.9% reduction was observed in membrane flux in 269 hours of
DCMD operation as seen in Figure 4.21. However, considerable change was not detected
in whole reduction process of membrane flux. It decreased slightly in long operation time.
The glucose fouling accumulated in membrane surface was discovered after 168 operation
hours at ~53% glucose concentration. The motivation of this phenomenon is not only the
concentration polarization at high concentrations but also due to the long exposure time.
Figure 4.22 Specific energy consumption in HF DCMD system with real glucose
liquid
Specific energy consumption that is presented in Figure 4.22 increased steadily with
increasing concentration. In HF DCMD, the energy consumption in real glucose liquid
(22.4 kW/kg at 61.3% glucose concentration) was higher than in synthetic liquid (11.69
kW/kg at 61.4%). The specific energy consumption is the ratio between utilized energy
0
20
40
60
80
0.0
0.4
0.8
1.2
0 50 100 150 200 250 300
Conce
ntr
atio
n (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Concentration
0
20
40
60
80
0
10
20
30
40
50
0 50 100 150 200 250 300
Conce
ntr
atio
n (
%)
Ener
gy c
onsu
mpio
n r
atio
(kW
/kg)
Time (h)
Flux Energy consumption ratio
79
and flux; therefore, the higher energy consumption in real glucose liquid case is explained
by the flux reduction. Actually, the energy consumption was relatively stable at 1.5 kWh.
At higher glucose concentration, the energy consumed for heating converted to energy for
pumping due to higher viscosity and homogeneous foulants in bulk feed.
Figure 4.23 Measured real glucose liquid concentration by DNS method in DCMD
configuration
4.3.1.3 Fouling analysis in real glucose liquid concentration on HF DCMD
Fouling analysis is very necessary in case of real glucose liquid on DCMD. The foulents
accumulated in membrane surface at high concentration was combination of different
fouling materials. The type of fouling and their contribution are presented in Table 4.20
and Figure 4.24.
Table 4.17 Different Type of Resistance in DCMD with Real Glucose Liquid
Resistance Value (106 m/s) Percentage (%)
Membrane 17.44 9.02
Boundary layers 21.68 10.76
Fouling
Recoverable 92.55 45.94
Organic Reversible 61.28 30.41
Inorganic reversible 5.37 1.99
Irreversible 3.16 1.21
Total 201.49 100
The fouling resistance accounted for ~ 79.5%, especially organic fouling played an
important role in membrane flux reduction.
25
35
45
55
65
0 5 10 15 20Glu
cose
conce
ntr
atio
n (
%)
Batch (12 hours)
Theoretical Glucose Concentration
Measured Glucose Concentration
80
Figure 4.24 Different types of resistances in DCMD with real glucose liquid
However, 31% organic fouling could be removed by washing with NaOH 2% and 2%
inorganic and fouling was cleaned by oxalic and citric acid solution.
Table 4.18 Recoverability from Organic and Biological Fouling of MD in Some
Research
Membrane
material Feed solution Cleaning method Recovery (%) Reference
PP Microbial biofilm
fouling in
seawater
NaOH at pH = 12
70% Ethanol for
disinfection, DI
water
≈100 (Krivorot et al.,
2011)
PTFE 61.3% glucose
liquid
NaOH 2%
0.1%w.t Oxalic,
0.8%w.t Citric
≈99 This study
4.3.2 Glucose liquid concentration on hollow fiber sweeping gas membrane
distillation (HF SGMD)
Real glucose liquid test was also conducted with HF SGMD to examine the performance of
membrane flux and energy consumption ratio. The experiment procedure was similar with
the test with HFDCMD with initial feed glucose concentration was 32%, the feed tank was
kept in 10L at the beginning of next batch by adding 10% glucose liquid. Glucose liquid
achieved 60% was the notification of the test completion. Operating condition was 50 oC
for feed temperature; 2.4 L/min for hot glucose liquid flow rate, gas flow rate was 25.5
L/min.
The HF SGMD system consumed around 60 hours of operation to concentrate real glucose
liquid from 32% to ~61.8%. The flux reduced slowly from 1.21 to 0.94 kg/m2.h that is
showed in Figure 4.25. As mentioned in Section 4.3.1, the TP effect is not serious in
membrane flux because TP is located in permeate side of SGMD. (Khayet et al., 2003)
concluded that membrane flux in SGMD is only dependent on feed temperature and air
Membrane
9%
Boundary
layers
11%
Recoverable
46%
Organic
31%Inoranic
2%
Irreversible
1%
Fouling
80%
81
flow rate. The solute concentration has effect on flux but only small reduction in flux is
found with higher solute concentration.
Figure 4.25 Permeate flux vs. feed real glucose liquid in HF SGMD system
The utilized energy for circulating feed liquid, heating and supplying gas in permeate are
grouped into energy consumption for SGMD configuration. Figure 4.26 describes the
transformation of the energy consumption ratio when increasing glucose concentration.
Figure 4.26 Specific energy consumption in HF SGMD system with real glucose liquid
The energy consumption ratio in SGMD configuration poorly increased as higher glucose
concentration. The lazily reduction of flux coupled with decreasing energy for heating due
to latent heat of high viscosity glucose liquid resulted in insignificant change of this energy
ratio. In SGMD, the specific energy consumption gently fluctuated at around 1.61 – 2.07
kW/kg.
0
10
20
30
40
50
60
70
0.0
0.4
0.8
1.2
1.6
0 10 20 30 40 50 60 70
Conce
ntr
atio
n (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Concentration
0
10
20
30
40
50
60
70
0.0
0.4
0.8
1.2
1.6
2.0
2.4
0 10 20 30 40 50 60 70
Conce
ntr
atio
n (
%)
Ener
gy c
onsu
mpti
on r
atio
(kW
/kg)
Time (h)
Flux Energy consumption ratio
82
Figure 4.27 Measured real glucose liquid concentration by DNS method in SGMD
configuration
The closeness of measured real glucose liquid concentration by DNS method and
theoretical value is showed in Figure 4.27.
4.3.2.1 Fouling analysis in HF SGMD with real glucose solution
Difference kinds of resistance were investigated with used membrane after concentration
glucose process in SGMD configuration. The cleaning process was similar with
concentrating real glucose liquid on HF DCMD. Particular resistances are presented in
Table 4.23 and Figure 4.28.
Table 4.19 Different Type of Resistance in SGMD with Real Glucose Liquid
Resistance Value (106 m/s) Percentage (%)
Membrane 17.44 37.64
Boundary layers 6.89 14.88
Fouling
Recoverable 14.62 31.56
Organic Reversible 7.19 15.51
Inorganic reversible 0.18 0.40
Irreversible 0 0
Total 46.32 100
The majority of resistance was raised in fouling resistance with accounted for 47% of total
resistance while membrane resistance also reached approximately 38%. Therefore, fouling
in SGMD with real glucose liquid just play a minor role in resistance.
25
35
45
55
65
0 1 2 3 4 5 6
Glu
cose
conce
ntr
atio
n(%
)
Batch (12 h)
Theoretical Glucose Concentration
Meaured Glucose Concentration
83
Figure 4.28 Different types of resistances in SGMD with real glucose liquid
4.4 The Comparison between DCMD and SGMD
The performance of DCMD and SGMD in term of specific energy consumption and
membrane flux was compared to specify the appropriate configuration for glucose liquid
feeding and divalent salt feeding test.
4.4.1 The comparison between HF SGMD and HF DCMD in TDS removal test
During the initial operation with the same condition, the flux for HF DCMD and HF
SGMD was relatively equal at 3.07 kg/m2.h as presented in Figure 4.29. In the DCMD, the
temperature difference between two sides (70 oC, 15 oC respectively) was greater than the
membrane SGMD (70 oC, 33 oC respectively). However, in SGMD configuration, gas
temperature contributes less effect to membrane flux. Ambient temperature was observed
to be good enough to utilizing as gas temperature (Khayet and Matsuura, 2011). The flux
of SGMD virtually insensitive with temperature of permeate gas that observed by Basini et
al. (1987). In addition, (Khayet and Matsuura, 2011) concluded that SGMD configuration
had smaller heat loss through the membrane compared to DCMD at the same conditions.
Figure 4.29 Flux comparison between HF SGMD and HF DCMD
Membrane
38%
Boundary
layers
15%
Recoverable
31.1 %
Organic
15.5 %
Inoranic
0.4 %Irreversible
0%
Fouling
47%
0.0
0.5
1.0
1.5
2.0
2.5
3.0
3.5
0 100 200 300 400 500 600 700 800
Flu
x (
kg/m
2.h
)
Concentration (g/L)
Flux of HF DCMD system Flux of HF SGMD system
84
However, the flux in HF SGMD system seemed to be stable at the initial concentrations
and fell down to 1.12 kg/m2.h at the saturated point, while, HF DCMD flux reduced more
significantly and it also reached 1.12 kg/m2.h at the end of the experiment. Therefore, the
membrane flux values at initial concentration and saturated point were similar between
DCMD and SGMD. In general of the whole process, as the salt concentration increases,
the HF DCMD flux decreased more sharply than flux in HF SGMD system. The difference
between two configurations was the changing trend of flux.
Figure 4.30 Energy ratio consumption comparison between HF SGMD and HF
DCMD
Figure 4.30 reflects the energy consumption ratio in HF DCMD system was markedly
higher compared with used energy ratio in HF SGMD system. This was due to the energy
was used for cooling water in DCMD system was much higher than energy utilized for
supplying gas in SGMD system. Moreover, the internal heat loss through thin membrane in
DCMD promoted more energy (in terms of heat conduction) and resulted in re-heating and
re-cooling the feed and permeate solution respectively.
4.4.2 The comparison between HF SGMD and HF DCMD in glucose
concentration test
The real glucose concentration experiments were conducted in both DCMD and SGMD
with the similar condition at 60 oC feed temperature, 2.4 L/min feed flow rate. The
difference was SGMD used air compressor to supply 25.5 L/min gas in permeate side,
while DCMD had to use cooler, heat exchanger and pump to provide cool water in
permeate.
0
2
4
6
8
10
0 100 200 300 400 500 600 700 800
En
ergy R
atio
consu
mpti
on
(kW
/Kg)
Concentration (g/L)
Energy consumption in HF DCMD system
Energy consumption in HF SGMD system
85
Figure 4.31 Flux comparison between HF SGMD and HF DCMD in glucose
concentration test
Figure 4.31 shows that the flux was reduced with glucose concentration increased in both
SGMD and DCMD. However, the SGMD flux was 1.5 times greater than the flux obtained
from DCMD at 32% glucose concentration, while at 61% glucose concentration, SGMD
flux was superior to DCMD flux with 7 times higher. Actually, the trends of flux show in
Figure 4.32 indicated that the flux was gradually decreased in both SGMD and DCMD.
There is no dramatically change of DCMD flux was found. There are two reasons are
given to explain the DCMD flux significantly smaller than SGMD flux when concentrating
glucose liquid from 32 to 61%. The first reason is higher thermal efficiency in SGMD than
in DCMD configuration or in other words, the DCMD heat conduction through membrane
material and gas filled in membrane pores is greater than in SGMD. Therefore, the latent
heat in DCMD is lower than in SGMD. (Khayet et al., 2003) admitted that conduction heat
accounted for 58.9 - 82.3% of total heat transfer in DCMD configuration, whereas, the heat
loss in SGMD is found only 9.5 - 28.6%. The second reason is long contact time between
membrane and high concentration of glucose. SGMD configuration only took 60 hours of
operation to increase the concentration of glucose solution from 32% to 61%, while
DCMD consumed 252 hours of operation. It created the favorable condition to promote the
fouling resistance in membrane surface (included organic fouling and biological fouling).
Figure 4.33 clearly shows that the dominant resistance in DCMD is fouling resistance,
whereas, it only plays a minor role in SGMD.
0
0.4
0.8
1.2
1.6
0
10
20
30
40
50
60
70
0 24 48 72 96 120 144 168 192 216 252 0 24 60
DCMD SGMD
Mem
bra
ne
Flu
x (
kg
/m2
.h)
Glu
cose
co
nce
ntr
ati
on
(%
)
Time (h)
Glucose concentrationFlux
86
Figure 4.32 Specific energy consumption comparison between HF SGMD and HF
DCMD in glucose concentration test
The energy consumption ratio in DCMD was significantly higher than in SGMD
configuration showed in Figure 4.32. On average, energy consumption ratio in DCMD
was ~9 times higher than the energy ratio consumed in SGMD. The root cause of this
phenomenon is the energy utilized for cool down the permeate water and pump cool water
in permeate side of DCMD is higher than the energy used for bring the sweeping gas in
permeate side of SGMD. In addition, the heat loss was higher in DCMD case thus the
supplement energy was required to re-heat the feed solution. The smaller DCMD flux is
also one of the reasons of lower this energy ratio in DCMD compared to SGMD.
Figure 4.33 Resistance comparison between HF SGMD and HF DCMD in glucose
concentration test
0
10
20
30
40
50
0
10
20
30
40
50
60
70
0 24 48 72 96 120 144 168 192 216 252 0 24 60
DCMD SGMD
En
erg
y r
ati
o (
kW
/kg
)
Glu
cose
co
nce
ntr
ati
on
(%
)
Time (h)
Glucose concentration
Energy consumption ratio
0%
20%
40%
60%
80%
100%
Real Glucose in
DCMD
Real Glucose in
SGMD
Fouling
Boundary layers
Membrane
87
Chapter 5
Conclusions and Recommendations
5.1 Conclusions
The operating condition of the system was selected by the pure water test with gas flow
rate of 25.5 L/min in SGMD and permeate water flow rate of 1.32 L/min in DCMD. The
feed temperature was chosen at 70 oC for TDS test and 50 oC for glucose test.
This study revealed the potential of high concentration of TDS removal on SGMD
configuration. The permeate flux was relatively stable at around 2.51 kg/m2.h when the salt
concentration increased from 40 to 450 g/L. The reason is low viscosity of salt solution has
less effect on the formation of boundary layers resistance. In addition, the temperature
polarization was observed as a favorable factor for membrane flux at high concentration.
The energy ratio consumed in this experiment was 1.07 kW/kg. High TDS removal was
also successful when operated in DCMD configuration. The membrane flux decreased
from 3.07 to 1.12 kg/m2.h when salt concentration increased from 40 to 450 g/L. The
energy consumption ratio rose from 3.32 to 8.44 kW/kg at higher salt concentration.
Fouling resistance did not play an important role in TDS test. The highest resistance
accounted for 45% that was localized in membrane resistance. Zero percent of irreversible
was found after cleaning the membrane with acidic solution.
The application of MD on concentrating glucose was fruitful in both DCMD and SGMD
configuration. DCMD configuration was able to concentration synthetic glucose liquid
from 10 up to 74% for 450 hours of operation with the flux reduced from 1.61 to 0.2
kg/m2.h. Real glucose liquid 32% was also conducted in DCMD. An 83.9% flux reduction
was observed after 269 hours operation due to fouling resistance accounted for 79.5% of
total resistance. The motivation of fouling accumulation is not only the concentration
polarization but also the long exposure time. Specific energy consumption increased
steadily with enhancement of glucose concentration from 8.3 to 22.4 kW/kg.
HF SGMD system consumed 60 hours to concentrate real glucose liquid from 32 to 61.8%.
The flux reduced slowly from 1.21 to 0.94 kg/m2.h. Temperature polarization effect is not
really resinous in SGMD flux when the glucose liquid increases because TP is located in
permeate side. Fouling resistance in SGMD did not play as a major role in resistance. It
accounted for 47%. Specific energy consumption gently fluctuated at around 1.61 – 2.07
kW/kg due to high latent heat of high viscosity glucose liquid.
In TDS removal test, there is no significantly difference of flux between DCMD and
SGMD configuration. However, the energy consumption ratio of DCMD system was
markedly higher compared with used energy ratio in SGMD system. The 7 times higher
flux, 4 times lower time consumption, 9 times lower energy consumption ratio compared
to DCMD make SGMD become an encouraging configuration in process intensification.
88
5.2 Recommendation for Further Study
The following recommendations are given for future research based on this study:
1. The fouling might be a problem by causing membrane flux reduction and damage
to membrane surface. The phenomenon of deposit formation can be mitigated by
creating secondary flow and increasing shear rate. The bubbles of inert gas with
appropriate size created in feed side in suitable duration is a promising strategy for
minimizing the fouling problem at high concentration of feed solution. An air
nozzle is implemented in the inlet of feed side to scatter air bubble.
2. In term of energy consumption, MD can use solar energy as the energy source.
However, solar energy should be converted to secondary energy (electricity)
because the heat temperature cannot be controlled by using directly solar energy, in
addition, the system should be operated 24 hours/day to minimize the fouling
formation.
3. The combination of membrane distillation and osmotic distillation is given to
improve the membrane flux. This is a modification of DCMD system. Instead of
cool water in the permeate side, the cool osmotic solution will be placed in
permeate side. The integrated of vapor pressure in MD and osmosis pressure in OD
promotes the membrane flux enhancement. However, the low temperature in
permeate side can result in crystallizing of osmotic solution. Therefore, the osmotic
solution with negative solubility is suggested.
4. The multistage membrane distillation with different feed concentration is
recommended. More than 1 MD systems are parallel connected. The next MD
system is operated with higher initial concentration compared with the initial
concentration of first MD system. This scenario can reduce the contact time of
membrane with high concentration of feed solution. The flux, energy consumption
and investment cost should be compared with single stage membrane distillation.
5. The simulation software such as Computational Fluid Dynamic is proposed to
observe the fluid flows in membrane module. The fouling location can be predicted
so that the appropriate solution can be given by changing the flow direction or
increasing feed flow rate.
6. LEP is very important in MD thus measurement of LEP before and after
conducting experiment.
89
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96
Appendix A
Membrane Module Dimension and Experimental setup Photo
97
Figure A1 Membrane module dimension (mm)
45
405325495
Ø6
Ø20
Ø44
Ø48 Ø56
98
Figure A2 Direct contact membrane distillation system
Figure A3 Sweeping gas membrane distillation system
Chiller Heat
exchanger
Heater =
Feed Tank
Control
box Permeate
tank
Control
box
Heater =
Feed Tank
Air flow
rate
controller
99
Appendix B
Experimental Results
100
Table B.1 Pure Water Flux of 1 µm, 100 µm of Thickness Flat Sheet Membrane
Time (h)
Weight of permeate (g) Permeate Flux (kg/m2.h)
1 2704.2
2 2827.5 6.9
3 2949.8 6.8
4 3070.5 6.7
5 3188.1 6.5
6 3313.9 7.0
7 3436.8 6.8
8 3557 6.7
9 3673.1 6.4
10 3785.4 6.2
11 3905.7 6.7
12 4010 5.8
13 4122.5 6.3
14 4239.5 6.5
Average Flux (kg/m2.h) 6.6
Table B.2 Salt Rejection Experiments of 1 µm, 100 µm of Thickness Flat Sheet
Membrane
Time
(h)
Weight of
permeate (g)
Permeate
Flux (kg/m2.h)
Feed conc.
(%)
Permeate con.
(%)
Rejection
(%)
1 1777.0 38
2 1875.1 5.4 38.21 0.00 99.99
3 1971.6 5.4 39.40 0.08 99.81
4 2065.9 5.2 41.44 0.37 99.12
5 2162.9 5.4 45.09 0.35 99.22
6 2244.09 4.5 44.99 0.85 98.11
7 2327.3 4.6 48.96 0.96 98.04
8 2390.5 3.5 51.54 1.90 96.31
9 2449.8 3.3 54.82 1.80 96.72
10 2504.0 3.0 56.27 3.10 94.49
11 2553.3 2.7 57.67 2.59 95.51
12 2636.8 4.6 59.45 8.32 86.01
13 2731.8 5.3 60.36 10.29 82.95
101
Table B.3 Pure Water Flux of 0.45 µm, 30 µm of Thickness Flat Sheet Membrane
Time (h) Weight of permeate (g) Permeate Flux (kg/m2.h)
1 2890.1
2 3407.6 28.8
3 3949.4 30.1
4 4473.6 29.1
5 5004.6 29.5
6 5564.6 31.1
7 6106.2 30.1
8 6657 30.6
9 7202.9 30.3
10 7754.9 30.7
11 8319.9 31.4
12 8872.5 30.7
13 9432.7 31.1
Average Flux (kg/m2.h) 30.3
Table B.4 Salt Rejection Experiments of 0.45 µm, 30 µm of Thickness Flat Sheet
Membrane
Time
(h)
Weight of
permeate (g)
Permeate Flux
(kg/m2.h)
Feed conc.
(%)
Permeate
con. (%)
Rejection
(%)
1 2737 38 0.00
2 2830 5.2 39.83 0.01 99.98
3 2890 3.3 41.14 0.01 99.98
4 2932.2 2.3 42.03 0.03 99.94
5 2975.1 2.4 43.07 0.06 99.87
6 3003.6 1.6 43.49 0.08 99.83
7 3034.1 1.7 44.61 0.11 99.76
8 3059.6 1.4 45.08 0.15 99.68
9 3082.9 1.3 45.88 0.18 99.61
10 3110 1.5 46.54 0.18 99.62
11 3133 1.3 47.20 0.25 99.48
12 3158.5 1.4 48.09 0.29 99.40
13 3179 1.1 48.70 0.32 99.35
102
Table B.5 Pure Water Flux of 0.45 µm, 80 µm of Thickness Flat Sheet Membrane
Time (h) Weight of permeate (g) Permeate Flux (kg/m2.h) 1 2799
2 7.7 7.7 3 8.0 8.0 4 8.5 8.5
5 8.1 8.1 6 7.9 7.9 7 7.6 7.6
8 8.5 8.5 9 7.3 7.3 10 7.8 7.8
11 8.7 8.7 12 7.8 7.8 13 7.9 7.9
Average Flux (kg/m2.h) 8
Table B.6 Salt Rejection Experiments of 0.45 µm, 80 µm of Thickness Flat Sheet
Membrane
Time
(h)
Weight of
permeate (g)
Permeate Flux
(kg/m2.h)
Feed conc.
(%)
Permeate
con. (%)
Rejection
(%)
1 2641.6 38
2 2717.6 4.2 39.50 0.00 99.99
3 2776 3.2 40.73 0.02 99.96
4 2818 2.3 41.68 0.14 99.67
5 2859.8 2.3 42.65 0.56 98.68
6 2904.7 2.5 43.76 1.24 97.16
7 2966.3 3.4 45.37 2.58 94.31
8 3041.5 4.2 47.4 4.22 91.10
9 3136.2 5.3 50.28 6.21 87.64
10 3268 7.3 51.45 8.50 83.47
11 3409.6 7.9 55.67 9.64 82.69
12 3555.1 8.1 60.29 11.48 80.97
13 3708.1 8.5 67.01 12.60 81.20
103
Table B.7 Pure water Test Experiment
Temperature
(oC)
Gas flow rate
(L/min)
Flux
(kg/m2.h)
Energy
(kW/h)
Energy ratio
(kW/kg)
50
16.9 0.52 0.36 2.70
19.6 0.78 0.40 2.00
22.5 1.05 0.42 1.58
25.5 1.05 0.46 1.71
28.5 1.05 0.57 2.14
60
16.9 1.31 0.47 1.39
19.6 1.57 0.51 1.27
22.5 1.83 0.62 1.34
25.5 1.83 0.67 1.43
28.5 1.57 0.66 1.64
70
16.9 1.83 0.63 1.35
19.6 2.09 0.70 1.31
22.5 2.61 0.71 1.07
25.5 3.14 0.87 1.09
28.5 2.70 0.91 1.32
Table B.8 Rejection Test
Time
(h)
Height Flux
(kg/m2.h)
Feed conc. (g/L) Permeate conc.
(× 𝟏𝟎−𝟑𝒈/𝑳)
Rejection
(%)
Gas flow rate 16.6 L/min
8.4
1 7.6 2.09 76.7 15.17 99.98
2 6.8 2.09 79.6 14.85 99.98
3 6.1 1.83 84.8 2.24 100.00
4 5.4 1.83 90.0 0.32 100.00
Gas flow rate 25.5 L/min
5.7
1 4.6 2.88 97.9 4.24 100.00
2 3.5 2.88 106.05 8.38 99.99
3 2.4 2.88 124.61 8.26 99.99
4 1.3 2.88 129.02 10.18 99.99
104
Table B.9 Testing the Capacity of HF SGMD with High Concentration Salt Solution
Time
(h)
Permeate Flux
(Kg/m2.h)
Na2SO4 solution concentration
(g/L)
Energy consumption
(kW/h)
1 43.6 3.07 0.5
2 47.9 3.07 0.5
3 52.5 2.79 0.5
4 58.3 2.79 0.5
5 65.4 2.79 0.5
6 74.4 2.79 0.5
7 86.5 2.79 0.5
8 103.1 2.79 0.5
8 100.9 2.79 0.5
9 110.4 2.79 0.5
10 121.8 2.79 0.5
11 135.8 2.79 0.5
12 153.4 2.79 0.6
13 176.3 2.79 0.6
14 207.2 2.79 0.45
15 246.1 2.51 0.5
15 239.1 2.79 0.5
16 260.5 2.79 0.55
17 286.0 2.79 0.5
18 313.7 2.51 0.5
19 347.3 2.51 0.5
20 389.0 2.51 0.5
21 435.4 2.23 0.45
22 494.4 2.23 0.45
22 474.7 2.23 0.5
23 508.6 2.23 0.5
24 547.7 2.23 0.5
25 593.3 2.23 0.5
26 647.3 2.23 0.5
27 712.0 2.23 0.4
28 749.5 1.12 0.3
105
Table B.10 Testing the HF SGMD with High Concentration Salt Solution in Real
Operation
Time
(h)
Permeate Flux
(Kg/m2.h)
Na2SO4 solution concentration
(g/L)
Energy consumption
(kW/h)
1 3.07 43.44 0.5
2 3.07 47.52 0.5
3 3.07 52.45 0.5
4 3.07 58.53 0.5
5 3.07 66.19 0.5
6 2.79 75.14 0.5
7 2.79 84.24 0.5
8 2.79 92.67 0.5
8 2.79 67.06 0.5
9 2.79 72.34 0.6
10 2.79 78.53 0.4
11 3.07 86.67 0.6
12 2.79 95.70 0.5
13 2.51 105.60 0.5
14 2.51 117.79 0.5
15 2.51 133.15 0.5
15 2.79 95.45 0.5
16 2.79 102.91 0.5
17 2.51 110.69 0.5
18 2.51 119.75 0.5
19 2.51 130.42 0.5
20 2.51 143.17 0.5
21 2.51 158.70 0.5
22 2.51 178.00 0.5
22 2.79 117.16 0.5
23 2.79 125.84 0.5
24 2.79 135.90 0.5
25 2.79 147.72 0.5
26 2.79 161.79 0.5
27 2.79 178.82 0.5
28 2.51 197.53 0.5
29 2.51 220.62 0.5
29 2.51 137.29 0.5
30 2.51 146.73 0.5
31 2.51 157.55 0.5
32 2.51 170.10 0.5
33 2.51 184.82 0.5
34 2.51 202.33 0.4
106
Time
(h)
Permeate Flux
(Kg/m2.h)
Na2SO4 solution concentration
(g/L)
Energy consumption
(kW/h)
35 2.51 223.50 0.6
36 2.79 252.91 0.5
36 2.23 149.12 0.5
37 2.23 158.22 0.6
38 2.79 171.30 0.5
39 2.51 185.06 0.5
40 3.07 205.22 0.7
41 2.51 225.29 0.5
42 2.51 249.72 0.5
43 2.51 280.09 0.5
44 2.51 159.40 0.5
51 2.44 306.22 0.5
51 2.51 178.56 0.5
52 2.23 189.46 0.5
53 2.23 201.79 0.6
55 2.51 236.38 0.5
58 2.42 314.17 0.5
58 2.23 291.60 0.5
59 2.23 311.54 0.5
60 2.23 334.40 0.5
61 2.23 360.89 0.45
62 1.95 387.77 0.45
63 1.95 418.97 0.5
64 1.95 455.63 0.4
65 1.39 486.00 0.3
107
Table B11 The Experiment of Sodium Sulfate Solution on HF SGMD with
Continuous Feeding
No. Flux
(kg/m2.h)
Calculated
Concentration (g/L)
Measured
Concentration (g/L)
Energy ratio
(Kw/kg)
1 40.0 38.772
2 3.07 43.6 0.97
3 3.07 47.9
0.97
4 2.79 52.5
1.07
5 2.79 58.3
1.07
6 2.79 65.4
1.07
7 2.79 74.4
1.07
8 2.79 86.5
1.07
9 2.79 103.1 99.496 1.07
10
93.0 73.88
11 2.79 100.9
1.07
12 2.79 110.4
1.07
13 2.79 121.8
1.07
14 2.79 135.8
1.07
15 2.79 153.4
1.21
16 2.79 176.3
1.21
17 2.79 207.2
1.00
18 2.51 246.1 201.112 1.19
19
221.0 196.2
20 2.79 239.1
1.07
21 2.79 260.5
1.14
22 2.79 286.0
1.07
23 2.51 313.7
1.19
24 2.51 347.3
1.19
25 2.51 389.0
1.19
26 2.23 435.4
1.25
27 2.23 494.4 392.7 1.25
28
445.0 359.8
29 2.23 474.7
1.33
30 2.23 508.6
1.33
31 2.23 547.7
1.33
32 2.23 593.3
1.33
33 2.23 647.3
1.33
34 2.23 712.0
1.16
35 1.12 749.5
1.97
36 2.23 445.0 457.792 1.33
108
Table B12 The Experiment of Sodium Sulfate Solution on HF SGMD with Practical
Feeding
No. Flux
(kg/m2.h)
Calculated
Concentration (g/L)
Measured
Concentration (g/L)
Energy ratio
(Kw/kg)
1 40 41.822
2 3.07 43.44 1.02
3 3.07 47.52 1.02
4 3.07 52.45 1.02
5 3.07 58.53 1.02
6 3.07 66.19 1.02
7 2.79 75.14 1.12
8 2.79 84.24 1.12
9 2.79 92.67 1.12
10 62.5
11 2.79 67.06 1.12
12 2.79 72.34 1.26
13 2.79 78.53 0.98
14 3.07 86.67 1.14
15 2.79 95.7 1.12
16 2.51 105.6 1.24
17 2.51 117.79 1.24
18 2.51 133.15 1.24
19 89
20 2.79 95.45 1.12
21 2.79 102.91 1.12
22 2.51 110.69 1.24
23 2.51 119.75 1.24
24 2.51 130.42 1.24
25 2.51 143.17 1.24
26 2.51 158.7 1.24
27 2.51 178 1.24
28 109.6
29 2.79 117.16 1.12
30 2.79 125.84 1.12
31 2.79 135.9 1.12
32 2.79 147.72 1.12
33 2.79 161.79 1.12
34 2.79 178.82 1.12
35 2.51 197.53 1.24
36 2.51 220.62 1.24
37 129
38 2.51 137.29 1.24
109
No. Flux
(kg/m2.h)
Calculated
Concentration (g/L)
Measured
Concentration (g/L)
Energy ratio
(Kw/kg)
39 2.51 146.73 1.24
40 2.51 157.55 1.24
41 2.51 170.1 1.24
42 2.51 184.82 1.24
43 2.51 202.33 1.16
44 2.51 223.5 1.32
45 2.79 252.91 1.12
46 141
47 2.23 149.12 1.4
48 2.23 158.22 1.57
49 2.79 171.3 1.12
50 2.51 185.06 1.24
51 3.07 205.22 1.27
52 2.51 225.29 1.24
53 2.51 249.72 1.24
54 2.51 280.09 245.856 1.24
55 2.51 159.4 1.3
62 2.44 306.22 1.3
63 167.7
64 2.51 178.56 1.24
65 2.23 189.46 1.57
66 2.23 201.79 1.4
67
68 2.51 236.38 1.24
69
70
71 2.42 314.17 1.29
72 274.06
73 2.23 291.6 1.4
74 2.23 311.54 1.4
75 2.23 334.4 1.4
76 2.23 360.89 1.31
77 1.95 387.77 1.5
78 1.95 418.97 1.6
79 1.95 455.63 1.4
80 1.39 486 452.684 1.81
110
Table B13 The Experiment of Appropriate Feeding Flow Rate Selection in HF
DCMD
No. Feed flow rate value
(L/m)
Flux (kg/m2.h) Energy (kWh)
1 0.74 2.61 2.57
2 1.06 2.61 2.62
3 1.32 2.88 2.67
4 1.69 2.88 2.70
5 2.1 2.88 2.68
6 2.51 2.88 2.74
7 2.68 2.88 2.77
Table B15 The Experiment of Sodium Sulfate Solution on HF DCMD with
Continuous Feeding
No. Flux
(kg/m2.h)
Calculated
Concentration (g/L)
Measured
Concentration (g/L)
Energy ratio
(Kw/kg)
40 50.572
1 3.07 43.6
3.32
2 3.07 47.9
3.32
3 3.07 53.1
3.32
4 3.07 59.6
3.32
5 3.07 67.8
3.32
6 3.07 78.8
3.32
7 2.79 92.4
3.52
8 2.79 111.7 96.256 3.52
9
66 76.64
10 3.07 72.3
3.45
11 3.07 79.8
3.45
12 2.79 88.2
3.8
13 2.79 98.6
3.8
14 2.79 111.8
3.8
15 2.79 129
3.66
16 2.79 152.4
3.66
17 2.79 186.3 164.352 3.52
18
117 110.012
19 2.79 127.1
3.73
20 2.51 137.8
4.06
21 2.51 150.4
4.06
22 2.51 165.6
4.06
23 2.51 184.3
4.06
24 2.51 207.6
4.06
25 2.51 237.8
4.06
26 2.51 278.2 176.776 4.06
111
No. Flux
(kg/m2.h)
Calculated
Concentration (g/L)
Measured
Concentration (g/L)
Energy ratio
(Kw/kg)
27
193 143
28 2.51 207.1
4.22
29 2.51 223.5
4.06
30 2.51 242.6
4.06
31 2.51 265.4
4.06
32 2.51 292.8
4.06
33 1.95 318.5
5.02
34 1.95 349
5.02
35 1.95 386 220.204 5.02
36
37 2.23 277 209.204 4.57
38 1.95 295
5.02
39 1.95 312.8
5.02
40 1.95 332.9
5.02
41 1.67 355.8
5.63
42 1.67 378
5.63
43 1.39 403.2
6.75
44 1.12 426.9 330.372 7.38
45
46 1.95 356 308.232 5.22
Table B16 The Experiment of Synthetic Glucose Liquid on HF DCMD
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Glucose
measurement (%)
0 9.7
1 10.0 1.61 2.3 5.60
2 10.3 1.61 2 4.87
3 10.6 1.61 2 4.87
4 11.0 1.61 2 4.87
5 11.3 1.61 2 4.87
6 11.7 1.61 2 4.87
7 12.1 1.21 2 6.49
8 12.4 1.21 2 6.49
9 12.8 1.21 2 6.49
10 13.2 1.21 2 6.49 9.6
11 10.8
12 11.1 1.61 2.1 5.11
13 11.5 1.61 2.3 5.60
14 11.8 1.61 1.9 4.63
15 12.3 1.61 2.1 5.11
16 12.7 1.61 2.3 5.60
17 13.0 1.21 2 6.49
18 13.4 1.21 2 6.49
19 13.8 1.21 2 6.49
112
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Glucose
measurement (%)
20 14.2 1.21 2 6.49
21 14.6 1.21 1 3.25 11.0 22 12.1
23 12.5 1.61 2 4.87
24 12.9 1.61 2 4.87
25 13.3 1.61 2 4.87
26 13.8 1.61 2 4.87
27 14.1 1.21 2 6.49
28 14.5 1.21 2 6.49
29 14.9 1.21 2 6.49
30 15.4 1.21 2 6.49 13.7 31 13.0
32 13.4 1.61 2 4.87
33 13.7 1.21 2 6.49
34 14.0 1.21 2 6.49
35 14.5 1.61 2 4.87
36 14.9 1.21 2.1 6.82
37 15.3 1.21 2 6.49
38 15.9 1.61 2 4.87
39 16.3 1.21 2 6.49
40 17.0 1.61 2 4.87
41 17.5 1.21 2 6.49 14.8
42 15.3
43 15.7 1.21 2 6.49
44 16.1 1.21 2 6.49
45 16.4 1.21 2 6.49
46 17.0 1.61 2 4.87
47 17.4 1.21 2 6.49
48 17.9 1.21 2 6.49
49 18.6 1.61 2 4.87
50 19.1 1.21 2 6.49
51 19.6 1.21 2 6.49
52 20.2 1.21 2 6.49 15.2 53 16.7
54 17.2 1.61 2 4.87
55 17.6 1.21 2 6.49
56 18.0 1.21 2 6.49
57 18.5 1.21 2 6.49
58 19.0 1.21 2 6.49
59 19.5 1.21 2 6.49
60 20.0 1.21 2 6.49
61 20.6 1.21 2 6.49
62 21.2 1.21 2 6.49
63 21.8 1.21 2 6.49 16.7 64 18.1
65 18.5 1.21 2 6.49
66 18.9 1.21 2 6.49
67 19.4 1.21 2 6.49
68 19.9 1.21 2 6.49
113
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Glucose
measurement (%)
69 20.4 1.21 2 6.49
70 21.1 1.61 2 4.87
71 21.7 1.21 2 6.49
72 22.3 1.21 2 6.49
73 23.0 1.21 2 6.49
74 23.7 1.21 2 6.49 18.6 75 19.6
76 20.0 1.21 2 6.49
77 20.5 1.21 2 6.49
78 21.0 1.21 2 6.49
79 21.6 1.21 2 6.49
80 22.1 1.21 2 6.49
81 22.7 1.21 2 6.49
82 23.4 1.21 2 6.49
83 24.0 1.21 2 6.49
84 24.7 1.21 2 6.49
85 25.5 1.21 2 6.49 20.7 86 21.0
89 22.6 1.21 2 6.49
92 24.4 1.21 2 6.49
96 26.8 1.01 2 7.79 25.1 97 22.2
100 23.9 1.21 2 6.49
103 25.9 1.21 2 6.49
108 28.5 1.01 2 7.79 25.8
107 23.6
110 25.5 1.21 2 6.49
113 27.6 1.21 2 6.49
117 30.5 1.01 2 7.79 29.1 118 25.1
121 26.8 1.21 2 6.49
124 28.6 1.07 2 7.31
128 31.3 1.01 2 7.79 29.8 129 26.4
132 28.1 1.07 2 7.31
135 30.0 1.07 2 7.31
139 32.9 1.01 1.75 6.82 30.0
140 27.6
143 29.5 1.07 2 7.31
146 31.3 0.94 2 8.35
150 34.1 0.91 1.75 7.58 30.9 151 28.6
154 30.6 1.07 2 7.31
157 32.5 0.94 1.9 7.93
161 35.4 0.91 1.825 7.90 31.9 162 29.6
165 31.5 1.07 2 7.31
168 33.4 0.94 2 8.35
172 36.2 0.91 1.75 7.58 36.3
114
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Glucose
measurement (%)
173 30.6
176 32.4 0.94 2 8.35
179 34.4 0.94 2 8.35
183 37.0 0.81 1.95 9.50 39.5
184 33.8
187 35.6 0.94 2 8.35
190 37.3 0.81 2 9.74
194 39.9 0.81 1.65 8.04 43.4 195 37.1
198 39.0 0.94 1.83 7.65
201 41.0 0.94 1.80 7.51
205 43.7 0.81 1.75 8.52 49.1 206 39.9
209 41.8 0.94 1.83 7.65
212 43.8 0.94 1.80 7.51
216 46.5 0.81 1.77 8.64 51.9
217 42.5
220 44.2 0.81 1.87 9.09
223 45.9 0.81 1.80 8.77
227 48.2 0.70 1.78 9.88 53.9 228 44.1
231 45.8 0.81 1.87 9.09
234 47.6 0.81 1.80 8.77
238 49.9 0.63 1.69 10.58 54.5 239 45.3
242 46.7 0.67 1.83 10.71
245 48.2 0.67 1.83 10.71
249 50.2 0.60 1.78 11.53 54.5
250 45.8
253 47.3 0.67 1.80 10.52
256 48.8 0.67 1.77 10.32
262 52.6 0.67 1.67 9.74 55.9 263 47.9
266 49.4 0.67 1.77 10.32
269 51.0 0.67 1.77 10.32
273 53.0 0.60 1.70 11.04 56.5 274 48.6
277 50.2 0.67 1.77 10.32
280 51.8 0.67 1.70 9.94
284 55.0 0.60 1.70 11.04 59.4
285 50.5
288 52.1 0.67 1.77 10.32
291 53.8 0.67 1.77 10.32
295 55.9 0.60 1.70 11.04 62.0 296 51.3
299 53.0 0.67 1.77 10.32
302 54.7 0.67 1.73 10.13
306 58.5 0.64 1.70 10.35 66.1 307 53.7
115
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Glucose
measurement (%)
310 55.0 0.67 1.60 9.35
313 56.8 0.67 1.60 9.35
314 57.3 0.60 1.50 9.74
317 58.6 0.50 1.43 11.17 66.4
318 56.3
321 57.3 0.54 1.69 12.34
328 60.9 0.50 1.57 12.27 66.9
329 58.5
332 59.9 0.54 1.67 12.18
335 61.4 0.54 1.60 11.69
339 63.4 0.50 1.60 12.47 68.4 340 60.8
343 62.0 0.54 1.67 12.18
346 63.2 0.40 1.60 15.58
350 64.9 0.37 1.48 15.80 72.1 351 62.4
354 63.6 0.40 1.67 16.23
357 64.7 0.40 1.60 15.58
361 66.0 0.30 1.58 20.45 71.7
362 63.4
364 64.2 0.40 1.60 15.58
367 65.4 0.40 1.60 15.58
372 67.5 0.36 1.55 17.14 68.4 373 64.9
376 66.1 0.40 1.63 15.91
379 67.4 0.40 1.67 16.23
383 68.7 0.30 1.58 20.45 69.1 384 66.1
387 67.3 0.40 1.93 18.83
390 68.7 0.40 1.97 19.16
394 70.5 0.30 1.60 20.65 70.5
395 67.8
398 69.2 0.40 1.60 15.58
401 70.6 0.40 1.53 14.94
405 72.0 0.30 1.48 19.16 71.8 406 69.3
409 70.7 0.40 1.50 14.61
416 74.2 0.31 1.40 17.53 72.2 417 71.4
420 72.9 0.40 1.43 13.96
423 73.9 0.27 1.47 21.43
427 74.9 0.20 1.40 27.27 73.01
116
Table B17 The Experiment of Real Glucose Liquid on HF DCMD
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Measured
Glucose (%)
0 32.0
3 33.6 0.81 1.87 9.09 6 35.3 0.81 1.70 8.28
9 36.9 0.67 1.67 9.74 12 38.7 0.67 1.63 9.55 36 13 33.5
16 34.9 0.81 1.72 8.38 25 39.7 0.64 1.65 10.19 38.8 26 34.9
29 36.3 0.67 2.16 12.62 32 37.9 0.67 2.08 12.16 35 39.2 0.54 1.96 14.32
38 40.6 0.54 1.93 14.12 42.9 39 36.1
42 37.6 0.67 1.77 10.32
48 40.1 0.54 1.68 12.30 51 41.6 0.54 1.57 11.44 49.1 52 37.2
55 38.7 0.67 1.23 7.21 58 40.0 0.54 1.67 12.18 61 41.3 0.54 1.70 12.42 44.5
62 38.3
68 40.8 0.54 1.42 10.35 71 42.2 0.54 1.33 9.74
74 43.7 0.54 1.97 14.37 44.9 75 39.4
78 40.7 0.54 1.57 11.44 81 42.0 0.54 1.57 11.44 84 43.4 0.54 1.83 13.39
87 44.9 0.54 1.23 9.01 46.2 88 40.5
91 40.0 0.66 1.47 8.77
100 43.8 0.44 1.47 13.19 46.6 101 40.6
104 41.9 0.54 1.60 11.69
107 43.2 0.54 1.60 11.69 110 44.3 0.40 1.30 12.66 113 45.5 0.40 1.43 13.96 48.1
114 42.2
123 45.7 0.45 1.50 13.15 126 46.9 0.40 1.40 13.64 47.2
127 43.7
130 45.1 0.54 1.57 11.44 133 46.2 0.40 1.57 15.26
136 47.3 0.40 1.57 15.26 139 48.6 0.40 1.43 13.96 49.2 140 45.3
117
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Measured
Glucose (%)
149 48.7 0.40 1.46 14.18 152 49.9 0.40 1.43 13.96 49.8
153 46.7
156 47.9 0.40 1.47 14.29 159 49.1 0.40 1.53 14.94
162 50.4 0.40 1.50 14.61 165 51.7 0.40 1.50 14.61 53.4 166 48.4
175 52.2 0.40 1.55 15.07 178 53.1 0.27 1.20 17.53 53.3 179 49.9
182 51.1 0.40 1.50 14.61 185 52.0 0.27 1.57 22.89 188 52.9 0.27 1.40 20.45
191 53.8 0.27 1.53 22.40 48.8 192 51.2
198 53.4 0.34 1.52 17.73
201 54.3 0.27 1.37 19.97 204 55.2 0.27 1.20 17.53 54.3 205 52.5
208 53.8 0.40 1.50 14.61 211 54.7 0.27 1.43 20.94 214 55.7 0.27 1.30 18.99
217 56.6 0.27 1.27 18.51 53.3 218 53.8
221 55.1 0.40 1.47 14.29
224 56.1 0.27 1.43 20.94 227 57.0 0.27 1.27 18.51 230 58.0 0.27 1.33 19.48 54.8
231 55.2
234 56.1 0.27 1.43 20.94 237 57.1 0.27 1.40 20.45
240 58.1 0.27 1.40 20.45 243 59.1 0.27 1.33 19.48 56.7 244 56.5
247 57.5 0.27 1.47 21.43 253 59.5 0.27 1.40 20.45 256 60.0 0.13 1.23 36.04 59.2
257 57.7
260 58.7 0.27 1.50 21.92 263 59.7 0.27 1.53 22.40
266 60.7 0.27 1.47 21.43 269 61.3 0.13 1.43 41.88 59.5
118
Table B18 The Experiment of Real Glucose Liquid on HF SGMD
Time Concentration
(%)
Flux
( kg/m2.h)
Energy
(kwh)
Energy ratio
( Kw/kg)
Measured
Glucose (%)
1 32.0
4 34.4 1.21 0.49 1.61 7 37.3 1.21 0.50 1.61
10 40.6 1.21 0.50 1.61 13 44.7 1.21 0.46 1.50 34.2 14 36.4
17 39.1 1.21 0.50 1.61 20 41.9 1.07 0.43 1.56 23 45.1 1.07 0.50 1.81
26 48.8 1.07 0.43 1.56 44.2 27 40.3
30 42.9 1.07 0.49 1.81
33 46.0 1.07 0.50 1.81 36 49.4 1.07 0.50 1.81 39 52.9 0.94 0.43 1.79 53.4
40 44.0
43 46.9 1.07 0.43 1.56 46 49.8 0.94 0.43 1.79
49 53.1 0.94 0.43 1.79 52 56.8 0.94 0.49 2.07 58.7 53 47.8
56 51.0 1.07 0.50 1.81 59 54.1 0.94 0.46 1.93 62 57.7 0.94 0.46 1.93
65 61.8 0.94 0.39 1.65 58.9
119
Appendix C
Details of Calculations
120
Appendix C.1 Calculation of Permeate Flux
The initial Permeate water flux at 60oC of 1 µm, thickness of 100 µm Flat sheet membrane
Initial weight (W1) = 2704.2 g
1-hour weight (W2) = 2827.5 g
Permeate Flux = 𝑊2−𝑊1
𝐴 =
2827.5−2704.2
1000 ×0.018= 6.85 kg/m2.h
Appendix C.2 Calculation of Rejection
Salt rejection of 1 µm, thickness of 100 µm Flat Sheet membrane
Feed initial salt = 38%
Feed final salt = 38.21%
Permeate = 0%
Rejection = 38.21−0
38.21× 100 = 99.99%
Appendix C.3 Calculation of LEP
LEP of 1 µm, thickness of 100 µm Flat Sheet membrane
𝐿𝐸𝑃 >−2𝐵𝛾𝑐𝑜𝑠𝜃
𝑟𝑚𝑎𝑥=
−2 . 1 . 72 . cos (112)
1 . 10−3= 54 (𝐾𝑃𝑎)
Appendix C.4 Calculation of membrane surface temperature
Table C1 General Information
Parameter Unit Feed Permeate
Bulk in oC 70 32
Bulk out oC 69 55
Viscosity µ Ns/m2 0.4072 × 10-3 2.02532 × 10-5
Specific heat capacity cp/ca kJ/ (kg.K) 4190.7 1009
Thermal conductivity k W/(mK) 0.66 0.0285
Density 𝜌 kg/m3 978.2 1.060
Heat of vaporization ∆H kJ/kg - 2454.4
Thermal conductivity of membrane
𝐾𝑚 = [𝜀
𝑘𝑔+
(1−𝜀)
𝑘𝑝]
−1
= [0.5
0.03+
(1−0.5)
0.29]
−1
= 0.0544 (𝑊
𝑚𝑘)
Prandlt number:
𝑃𝑟 =𝜇𝑐𝑝
𝑘=
0.4072 × 10−3 × 4190.7
0.66= 3.5855
Hydraulic diameter:
121
Available space area = Amodule – Amembrane
= [𝜋×0.0482
4] − 100 × (
𝜋×0.002032
4) = 1.4859 × 10-3 (m2)
While A = π (𝑅2 − 𝑟2)
R = 0.048 m r = 0.0428 (m)
Dh = R × 2 – r × 2 = 0.048 × 2 – 0.0428 × 2 = 0.0104 (m)
Reynolds number
𝑅𝑒 =𝑣 × 𝐷ℎ × 𝜌
µ=
0.0269 × 0.0104 × 978.2
0.4072= 672.0561 (
𝑚
𝑠) < 800
Nusselt number
𝑁𝑢 = 1.86 × (𝑅𝑒 × 𝑃𝑟 ×𝐷ℎ
𝐿)
13
= 1.86 × (672.0561 × 2.5855 ×0.0104
0.4)
1
3
= 6.6245
Heat transfer coefficient
ℎ𝑓 =𝑁𝑢×𝑘
𝐷ℎ= 420.401 W/(m2.h)
Table C2 Membrane Surface Temperature Calculation
Permeate side Feed side
Using Equation 2.41
Q = 2718.12 W/m2
Using Equation 2.40
Ha = 163.62
Using Equation 2.39
Tmp = 60 oC
Assume Tm,f = 65 oC and Tm,p = 60 oC
Using Equation 4.5 Hv = 2612.6 kJ/kg
Equation 2.36 is used to evaluate Tmf = 67.6 oC
Reapeating the same procedure
Hv = 2136.1733 kJ/kg
Recalculating Tm,f = 67.7 oC constant
Therefore, membrane surface temperature:
Tm,f = 67.7 oC
Tm,p = 60 oC
Appendix C.5 Calculation of temperature polarization coefficient
𝑇𝑃𝐶 =𝑇𝑚,𝑓 − 𝑇𝑚,𝑝
𝑇𝑏,𝑓 − 𝑇𝑝,𝑏=
67.7 − 60
69.5 − 43.5= 0.3
Appendix C.6 Calculation of membrane distillation coefficient and resistance
122
Table C3 Temperature Information of Salt test in SGMD configuration
Temperature (oC) Flux
Kg/m2.h Bulk Feed Membrane feed Bulk permeate Membrane permeate
69.5 67.7 43.5 60 3.14
Feed side:
𝑃𝑤,𝑓 = 𝑒𝑥𝑝 (23.1964 −3816.44
𝑇 − 46.13) = 𝑒𝑥𝑝 (23.1964 −
3816.44
340.7 − 46.13) = 28013.2 𝑃𝑎
Permeate side:
𝑃 = 𝑒𝑥𝑝 (23.1964 −3816.44
𝑇 − 46.13) = 𝑒𝑥𝑝 (23.1964 −
3816.44
333 − 46.13) = 19784.9 𝑃𝑎
𝜔 = 𝜔𝑖𝑛 +𝐽𝑤 × 𝐴
𝑚𝑎= 0.014 +
3.14
3600×
0.255
4.505. 10−4= 0.5077
Coefficient:
𝑃𝑤𝑝 =𝜔 × 𝑃
𝜔 + 0.622=
0.5077 × 19784.9
0.5077 + 0.622= 8891.7 𝑃𝑎
Resistance:
𝐵𝑤 =𝐽
𝑃𝑤𝑓 − 𝑃𝑤,𝑝=
3.14 ÷ 3600
28013.2 − 8891.7= 4.56 × 10−8 𝑠/𝑚
𝑅𝑤 = 𝐵𝑤−1 = 21.9 × 106𝑚/𝑠
Appendix C.6 Calculation of fouling resistance
Table C4 Temperature Information of Salt test in SGMD Configuration
Temperature (oC) Initial Flux
Kg/m2.h
Final flux
Kg/m2.h Bulk Feed Mem. feed Bulk permeate Mem. permeate
69.5 67.7 43.5 60 3.07 1.95
Total fouling resistance (Rt)
𝑅𝑡 =𝑃𝑓 − 𝑃𝑝
𝐽=
30306.5 − 2987.9
1.95 ÷ 3600= 50.43 × 106 𝑚/𝑠
Feed boundary layer resistance (Rb,f)
𝑅𝑏,𝑓 =𝑃𝑏,𝑓 − 𝑃𝑚,𝑓
𝐽=
30306.5 − 28013.2
1.95 ÷ 3600= 4.23 × 106 𝑚/𝑠
Permeate boundary layer resistance (Rb,p)
123
𝑅𝑏,𝑝 =𝑃𝑏,𝑝 − 𝑃𝑚,𝑝
𝐽=
28013.2 − 6729.3
1.95 ÷ 3600= 6.91 × 106 𝑚/𝑠
Boundary layers resistance (Rb)
𝑅𝑏 = 𝑅𝑏,𝑓 + 𝑅𝑏,𝑝 = 11.14 × 106 𝑚/𝑠
Fouling resistance (Rf)
𝑅𝑓 = 𝑅𝑡 − (𝑅𝑏 + 𝑅𝑚) = 16.7 × 106 𝑚/𝑠
Recoverable fouling (Rr): Rf,r = 3.89×106 was calculated with flux = 1.95 kg/m2.h after
rinsing with DI water 30 min
𝑅𝑟 = 𝑅𝑓 − 𝑅𝑓,𝑟 = 16.7 × 106 − 3.89 × 106 = 12.85 × 106 𝑚/𝑠
Reversible fouling (Rre): with flux = 3.07 kg/m2.h after cleaning with acidic solution and
DI water
𝑅𝑟𝑒 = 3.89 × 106 − 0 = 3.89 × 106 𝑚/𝑠
Irreversible fouling (Rir)
𝑅𝑖𝑟 = 𝑅𝑓 − (𝑅𝑟 − 𝑅𝑟𝑒) = 0
124
Appendix D
Membrane Fouling
125
New
membrane Fouled membrane
After rinsing
with DI water
After cleaning
with chemical
Divalent
Salt
DCMD
Glucose
SGMD
Glucose
126
Appendix E
Experimental Activities
127
Appendix E1 Direct contact membrane distillation system control
Preparation
Startup the heater and cooler;
Valve 1, 3 close, valve 2 open to circulate permeate water through heat exchanger;
Valve 4 is used to control the feed flow rate.
Operation
Valve 3, 1 are opened, valve 2 is closed when temperature in feed and permeate
side reached the desired temperature;
Record the initial height of feed tank, energy consumption, temperatures in control
box.
Appendix E3 Analysis activities
Figure E1 The color transformation of the glucose curve
1
2
3
4
Feed out
Permeate
out
Feed in
Permeate
in
128
Appendix F
Checking the Flat Sheet Membrane
129
The Flat sheet DCMD unit was to evaluate the best membrane among 3 types of flat sheet
membranes. Three different pore sizes and thickness membranes used in this stage were
1µm, 100µm of thickness; 0.45µm, 30 µm of thickness and 0.45 µm, 80 µm of thickness.
The flat sheet is selected because it has higher flux and less fouling problem than hollow
fiber.
The condition for testing experiment is showed on Table 4.1. The information need to be
noted is the flow rate of feed and permeate side have to equal because when the different
pressure between two side of membrane excess LEP value, liquid will penetrate through
hydrophobic membrane and the non-volatile compounds also pass to permeate side.
Therefore the criteria for selecting a suitable membrane are a high flux membrane, besides
that the hydrophobic characteristic of membrane must be secured.
Table F.1 Checking the Membranes Condition
Membrane
Experiment Condition Parameter
measurement Pore
size Thickness
0.45 µm 30 µm
Pure water test
Rejection test
Feed temperature: 60oC
Permeate temperature: 10oC
Time: 12 hours/ test
Salt concentration: 380g/L
Flow rate: 1 L/min
Membrane Flux
Rejection
80 µm
0.1 µm 100 µm
F.1 Membrane pore size 1 µm, thickness 100 µm
The pure flux of the membrane is quite low; it is around 6.7 kg/m2.h. Pure water membrane
flux values fluctuate slightly, the cause may be due to the ability to maintain the
temperature as set on the feeding and the permeate side is not very good.
The result from figure 1 showed that the rejection decreasing from 99.99 % to 82.95 %
after 13 hours of operation. The flux decreasing at the first hours because of the increasing
concentration of salt in feed side and the lowest value was around 2.7 kg/m2.h, then the
flux increasing while the rejection decreasing, this phenomenon prove that the rejection
capacity of membrane is reduced because of higher concentration of salt and there are
some salt penetrate through membrane, in other words, the pressure difference between
two sides of membrane excess the LEP value. Besides that, as I observed, there is some
crystal salt accumulated in permeate side of membrane module after 13 hours operation
with high salt concentration.
130
Figure F.1 Membrane Flux of FS DCMD Pore Size 1 µm, Thickness 100 µm
The membrane pore size 1 µm, thickness 100 µm has ability to concentrate divalent salt
(Na2SO4) from 380 mg/L to around 603 mg/L in 10 hours. After 10 hours of operation
pressure difference between two side of membrane excess LEP value therefore from this
point toward the flux is higher and rejection start to reduce. LEP value can be calculated by
equation 2.1, LEP for this type of membrane is 54 KPa.
After 10 hours of operation, the mass balance of the system can be demonstrated in
following figure:
Feed out
603.4 g
Feed in
760 g
Permeate
102.9 g
Figure F.2 Mass Balance of Salt in FS DCMD with Pore Size 1 µm, Thickness 100 µm
after 12 hours Operation
The mass of salt in the feed side out was lower than the salt in feed side in because some of
salt penetrated to the feed side. The amount of salt in feed side in was nearly equal to the
total salt in feed side out and permeate side because some salt were also accumulated in
membrane surface.
0
20
40
60
80
100
0
1
2
3
4
5
6
7
8
1 2 3 4 5 6 7 8 9 10 11 12 13
Rej
ecti
on (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Pure water Flux Rejection
131
F.2 Membrane pore size 0.45 µm, thickness 30 µm
The pure water flux of membrane pore size 0.45 µm, thickness 30 µm is much higher than
the flux of membrane pore size 1 µm and thickness 100 µm. The flux is highest at 31.4
kg/m2.h.
Figure F.3 Membrane Flux of FS DCMD Pore Size 0.45 µm, Thickness 30 µm
The high concentration of salt solution had a greatly effect on membrane flux. As the result
from figure 3, the flux of membrane was remarkably discrepancy when the feeding was
salt solution with concentration 380 mg/L (2.67 mol/L), the flux value was around 5
kg/m2.h. The permeate flux with membrane pores size 0.22 µm , 35 µm thickness with
feeding solution is 2 mol/L of divalent salt solution is also very low, at around less than 5
kg/m2.h as (Guan et al., 2015) also proved.
The rejection capacity of membrane pore size 0.45 µm, thickness 30 µm is very high
(around 99%). Almost salt was not penetrate through this kind of membrane. The
membrane can concentrate divalent salt from 380 mg/L to 487 mg/L in 12 hours.
The LEP of membrane is inversely proportional to membrane pore size; therefore, the
membrane pore size 0.45 µm, thickness 30 µm has higher LEP (120 KPa) value in
compare to membrane pore size 1 µm, thickness 100 µm.
After 10 hours of operation, the mass balance of the system can be demonstrated in
following figure:
0
20
40
60
80
100
0
5
10
15
20
25
30
35
1 2 3 4 5 6 7 8 9 10 11 12
Rej
ecti
on (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Pure water Flux Rejection
132
Feed out
758.8 g
Feed in
760 g
Permeate
1.4 g
Figure F.4 Mass Balance of Salt in FS DCMD with Pore Size 0.45 µm, Thickness 30
µm after 12 hours Operation
F.3 Membrane pore size 0.45 µm, thickness 80 µm
The membrane has quite low pure water flux. It fluctuates slightly with an average value of
about8 kg/m2.h. The thickness is an important property that has significantly effect to
membrane flux; the membrane with thicker thickness has lower flux (Shirazi et al., 2014)
In addition, the rejection capacity of the membrane was very low; the rejection reached
99% in the first three hours, then the ability of the membrane to refuse salt decreased to
84.5% after 12 hours of operation, simultaneously, the membrane flux keep increasing
after first 3 hours because of the loss of membrane selectivity. Therefore, this kind of
membrane just able to work at salt concentration is less than 407 g/L of divalent salt.
Figure F.5 Membrane Flux of FS DCMD Pore Size 0.45 µm, Thickness 80 µm
0
20
40
60
80
100
0
2
4
6
8
10
12
1 2 3 4 5 6 7 8 9 10 11 12 13
Rej
ecti
on (
%)
Flu
x (
kg/m
2.h
)
Time (h)
Flux Pure water Flux Rejection
133
Feed out
625.4 g
Feed in
760 g
Permeate
134.4 g
Figure F.6 Mass Balance of Salt in FS DCMD with Pore Size 0.45 µm, Thickness 80
µm after 12 hours Operation
F.4 The comparison between the three different types of membranes
According to figure 7, the membrane pore size 0.45 µm, thickness 30 µm has highest pure
water flux (30.3 kg/m2.h) compared to two remaining membranes. The reason is that the
thickness may has a significantly effect to membrane flux.
Figure F.7 The Pure Water Flux Comparison of 3 Type of Flat Sheet Membranes
0
5
10
15
20
25
30
35
1 2 3 4 5 6 7 8 9 10 11 12
Flu
x (
kg/m
2.h
)
Time (h)
Pore size 1 micron, thickness 100 micron
Pore size 0.45 micron, thickness 30 micron
Pore size 0.45 micron, thickness 80 micron
134
Figure F.8 The Salt Rejection Capacity Comparison of 3 Type of Flat Sheet
Membranes
As the results showed in figure 8, membrane pore size 0.45 µm, thickness 30 µm has
highest capacity of salt rejection. The membrane was able to work at very high
concentration of divalent salt (487 mg/L) in 12 hours with the rejection capacity reached
more than 99%. However, at higher concentration of salt, the membrane flux reduced.
Membrane pore size 1 µm, thickness 100 µm has higher flux when it works at high
concentration of salt than membrane pore size 0.45 µm, thickness 30 µm, that is the reason
for the membrane can reaches very high concentration of salt (603 mg/L) after just 10
hours of operation.
0 380 600487 Salt Concentration
(mg/L)
Pore size 1 micron,
Thickness 100 micron
Pore size 0.45 micron,
Thickness 30 micron
Flux: 31.4 kg/m2.h Flux: 2 kg/m2.h
Flux: 4.5 kg/m2.hFlux: 6.6 kg/m2.h
Figure F.9 The Flux Comparison between Membrane Pore Size 0.45 µm, Thickness
30 µm and Membrane Pore Size 1 µm, Thickness 100 µm
Therefore, selecting the suitable membrane should be based on concentration of feeding
solution. There are 2 types of membranes that can be selected as membrane pore size 0.45
µm, thickness 30 µm and membrane pore size 1 µm, thickness 100 µm because they have
high flux and rejection capacity. However, if the feeding solution has low concentration,
membrane pore size 0.45 µm, thickness 30 µm should be used because it has higher flux at
low concentration of feeding solution, in contrast, if the feeding solution has high
concentration (around 300 to 600 mg/L), it is better to choose membrane pore size 1 µm,
thickness 100 µm.
0
20
40
60
80
100
1 2 3 4 5 6 7 8 9 10 11 12
Flu
x (
kg/m
2.h
)
Time (h)
Pore size 1 micron, thickness 100 micron
Pore size 0.45 micron, thickness 30 micron
Pore size 0.45 micron, thickness 80 micron