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THORSEN 2006 Nanofiltration in Drinking Water Treatment
Techneau, D5.3.4B
December 2006
Nanofiltration in drinking water
treatment
Literature Review
Techneau, 11.
December 2006
Nanofiltration in drinking water
treatment
Literature Review
© 2006 TECHNEAU
TECHNEAU is an Integrated Project Funded by the European Commission under the Sixth Framework
Programme, Sustainable Development, Global Change and Ecosystems Thematic Priority Area
(contractnumber 018320). All rights reserved. No part of this book may be reproduced, stored in a database
or retrieval system, or published, in any form or in any way, electronically, mechanically, by print,
photoprint, microfilm or any other means without prior written permission from the publisher
Colophon
Title
Nanofiltration in drinking water treatment
Author(s)
Thor Thorsen, Harald Fløgstad
Quality Assurance
By Farhad Salehi
Deliverable number
D 5.3.4B
This report is:
PU = Public
Contents
Contents
1
1
Introduction
3
1.1
Classification of membrane filtration
3
1.2
Requirements in drinking water treatment
6
1.3
Properties of direct nanofiltration (NF)
7
2
Filtration and fouling mechanisms
10
2.1
Particle characterisation
10
2.2
Filtration and fouling
12
2.3
Causes of residual fouling in practical filtration
14
2.4
The significance of the membrane type
16
2.5
Chemical factors in NOM fouling
17
3
Fouling control
3.1
General experiences
20
3.2
3.2.1
3.2.2
Laboratory and pilot experiments
Spiral wound membranes versus capillary membranes
Spiral wound membranes with different prefilters
21
21
22
3.3
Fouling and rejection at different plant recovery
24
4
Experience in NOM removal applications
4.1
Norwegian experiences with spiralwound NF membranes
25
4.2
Configurations with tubular membranes
29
5
Groundwater and softening applications
5.1
5.1.1
5.1.2
Hardness
Types of hard water
Traditional softening methods
32
32
33
5.2
5.2.1
5.2.2
Scaling
Scale control
Antiscalants
33
34
35
5.3
5.3.1
5.3.2
5.3.3
5.3.4
Case studies
Florida
Mainz, Germany
Spain
England
36
36
38
39
41
5.4
Waste disposal
43
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20
25
31
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6
Summary
45
7
Conclusions
48
References
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1 Introduction
1.1 Classification of membrane filtration
The term “membrane filtration” describes a family of separation methods.
The basic principle is to use semi-permeable membranes to separate fluids,
gases, particles and/or solutes. Membranes are usually shaped as a thin film,
which allows transport of some materials, but not all. For separations from
the water phase the membrane is water-permeable, but less permeable to
solutes and other particles depending on their size and to some degree other
properties. All living organisms rely on natural membrane selective transport
of solutes in to and out of biological cells. Membranes are the active barriers
in organs like kidneys and the stomach. Although membrane filtration is a
relatively new family of methods for technical filtration, the principles of
most methods have been known for some time.
Semi-permeable membranes have pores in the range 0.5 nm to 5 µm. Figure 1
illustrates which compounds can be separated. Most filter membranes are
produced with physical/chemical methods where the pores are formed by
physical and chemical processes. An important property that characterises the
individual membrane methods is the driving force behind the separation.
Some methods are summarised in Table 1, showing their driving force,
membrane structure and the approximate time of introduction for technical
filtration. It can be seen that the driving force is different, and so is the design
of the technical filter equipment.
One of the methods is crossflow filtration, or tangential flow filtration, which
is focussed on in this report. Of the methods in Table 1, crossflow filtration
has the widest application and a major application is in drinking water
treatment. Other methods are used in industrial separations, although
electrodialysis and membrane distillation may potentially be used for
drinking water treatment as well.
Size scale
0 .001
0.1
1.0
1
10
1 00
Polysaccarides
Simple
and proteins
organics
Humic substances
Viruses
1 0 00
Inorg
ions
0 .0 1
Colloids
nm
10
10 0
1 000 µ m
0 .0 1
0.1
1 .0 m m
Bacteria
F ilte r ch a n n el
d im ensio ns
Visible particles
Figure 1. An overview of the relevant dimensions in membrane filtration. The pores of filtration
membranes range from about 0.5 to more than 1000 nm.
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Table 1. Properties of various methods for membrane filtration [1].
Method
Driving force 1)
Membrane Permeation
Introduced 2)
Dialysis
Concentration
Porous
Solutes
1950
Electrodialysis
Electrical
Porous
Ions
1955
Crossflow filtration
Pressure/concentration Porous
Water
1960
Pervaporation
Partial pressure
Dense
Liquid
1982
Porous
Liquid
1981
Membrane distillation Partial pressure
1) Potential
2)
and gradient that enforce permeation.
For technical filtration
In crossflow filtration the bulk flow in the filter is along the membrane
surface and perpendicular to the direction of filtration. Water permeates the
membrane as the feed flow passes by. The transmembrane pressure drives
water transport through the membrane and permeable particles are
transported through the membrane, often driven by a concentration gradient.
Both diffusion processes and convective flows are essential in the process.
The basic principle is illustrated in Figure 2.
The membrane pore size is a main factor determining whether a solute will
pass the membrane. A classification of crossflow filtration is given in Table 2,
showing that the pore sizes cover a very wide range from less than one
nanometer to more than one micrometer. The definition of transition values
between the methods varies somewhat in the literature. Dalton is a common
designation of molecular weight in membrane filtration and expressed in
g/mole.
Consentrate channel
Particles and m olecules
Pum p
FEED
CONCENTRATE
Mem bran section
PERMEATE
Figure 2. The basic principle of crossflow filtration.
The basic equations that describe the filtration are relatively simple:
Water transport (flux):
Jw = A · (P – π)
[L/m2h]
Solute transport:
Js = B · (cC – cP)
[g/m2h]
Rejection:
R = (cC – cP) /cC
[% or fraction]
(1)
(2)
(3)
where A is the water permeability [L/(m2 h bar)], P the transmembrane
pressure [bar], π the osmotic pressure [bar], B the solute permeability
[L/(m2 h)], cP the concentration in permeate [g/L], and cC the concentration in
concentrate [g/L].
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Osmotic pressure is a colligative property of solutions. It means that a
minimum transmembrane pressure must be applied for water transport
through the membrane to occur. The osmotic pressure π is proportional to the
total difference in solute particle concentration across the membrane. In
desalination of seawater its value is about 26 bars, but in surface water
treatment the concentrations are too low to give significant osmotic pressure.
In Equation 2 the solute transport is driven by a concentration difference. R is
the local particle rejection of the membrane and is an important parameter. It
should not be confused with treatment efficiency in a technical membrane
plant, which is similar, but based on plant feed and total permeate quality
instead of local concentration values.
Table 2. Properties of individual crossflow methods [2]
Method and
abbreviation
Reverse osmosis, RO
Pore size
Molecular
nm
weight cutoff1)
Pressure
Permeation
bar
< 0.6
< 500
30 –70
Water
Nanofiltration, NF
0.6 – 5
500 – 2000 Da
10 – 40
Water, low molecular solutes
Ultrafiltration, UF
5 – 50
2 – 500kDa
0.5 – 10
As above plus macromolecules
Microfiltration, MF
50 – 5000
> 500 kDa
0.5 – 2
As above plus colloids
1)
Molecular weight (Dalton) cutoff of the membrane, where solutes of this weight are rejected by 90%
Membranes are usually made from synthetic organic polymers and the
thickness is in the order of 0.2 mm for sheet membranes. The physical shape
of the membrane is designed to fit in suitable “modules”. A number of
membrane module types are made, using sheet as well as hollow fibres,
capillary or tubular membranes. Capillary membranes have achieved a
certain foothold for drinking water, mainly because they can be backflushed
to remove deposits. Hollow fine fibres are common in desalination of
seawater. Spiral modules are popular for drinking water because of their low
cost and moderate fouling tendency. Sketches of spiral and capillary module
types are shown in Figure 3. Plate and frame and tubular systems are bulky
and expensive, but are still used in a few smaller plants for drinking water
treatment.
A prefilter is an essential part in a membrane plant in order to prevent that
particles larger than the size of the narrow channels between the membranes,
commonly 0.7–2 mm, enter the modules. Still some accumulation of matter on
the membrane surface takes place and eventually reduces the flux and the
capacity of the plant. This phenomenon is referred to as fouling. Avoiding
and controlling fouling is the most important challenge for successful
membrane filtration.
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Figure 3. The spiral (left) and capillary modules (right) are the most actual types drinking water
treatment.
1.2 Requirements in drinking water treatment
Membrane filtration was introduced in drinking water treatment in the 1950s,
mainly for desalination of seawater, brackish water and groundwater.
Membrane filtration features the unique property that a membrane can be
chosen that removes just the components that is needed from the actual raw
water. Such components typically are [2],[3]:
•
•
•
•
•
•
•
Inorganic or organic salts
Metals
NOM
Biodegradable organics
Disinfection by-products
Turbidity and particles
Infectious species (bacteria, virus, parasites)
Traditionally the most used applications in drinking water treatment have
been:
•
•
•
Desalination of seawater or brackish water
Removal of hardness, typically from groundwater (softening)
Turbidity and bacteria removal
Since the late eighties an increase in the number of plants used for treatment
of surface water has also been seen. These are used for treatment purposes
like removal of infectious species, turbidity, hardness, micropollutants, NOM
and taste and odour. Different treatment needs membranes of different pore
diameters, as shown in Table 3. In this report the focus is mainly on natural
organic matter.
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Taste and odour is a special case for membranes in that the chemical nature of
such compounds is highly variable, from relatively large organic molecules to
low-molecular compounds. Often the source of taste and odour is volatile
compounds that are typically low- molecular, and in these cases RO may be
needed. In that case the use of activated carbon or ozone as a secondary
treatment may by the best solution.
In the Nordic countries colour removal is an actual treatment. This means
removal of natural organic matter, NOM, typically humic substances. Colour
is just one character of NOM. Countries where treatment of coloured surface
water is actual are boggy regions in cold climates with limited ground water
resources. This is typical for some areas in Northern North America, Great
Britain, Scandinavia and Russia. Coloured surface water in these areas
typically is soft and has high concentration of soluble NOM with significant
colour. Compared with other areas where NOM is a concern, the raw water
has a higher concentration of humic substances and less salinity.
Membranes for desalination of seawater and brackish water have pores
around 0.5 nm. Removal of larger particles like bacteria and in turbidity, calls
for much more open membranes with pores of 10 nm or larger. If virus
removal is not an issue, pores above 100 nm are also applicable. This is in the
crossflow microfiltration range. For removal of humic substances colour is the
most relevant parameter and pore sizes in the range 1 – 5 nm are most
relevant. This is summarised in Table 3 [2].
Table 3. Properties of various drinking water plants [2], see Table 2
Parameter
Process
Raw water types
Seawater Groundwater High col. Medium col. Bacteria, SS
DesalinaSoftening
Colour removal
Particle removal,
tion
disinfection
Plant:
- Operating pressure 50 – 60 bar
< 0.5 nm
- Pore diameter
RO
- Membrane methods
Removal efficiency:
> 96 %
- Colour (in NOM)
100 %
- Susp. matter
100 %
- Salts
100 %
- Bacteria
100 %
- Virus
10 – 20 bar
0.5 –1 nm
RO, NF
4 – 8 bar
1 – 2 nm
NF
2 – 5 bar
2 – 5 nm
NF, UF
0.5 – 2 bar
5 – 200 nm
UF, CMF
> 96 %
100 %
100 %
100 %
100 %
90 – 95 %
100 %
30 %
100 %
100 %
80 – 90 %
100 %
< 20 %
100 %
100 %
10 – 80 %
90 – 100 %
<5%
100 %
10 - 100 %
1.3 Properties of direct nanofiltration (NF)
NOM is a polydisperse mixture of individual particles in natural water
originating from degraded and partly re-synthesised plant residuals. Natural
concentrations of natural organics are low, usually less than 20 mg/L, but
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their removal is important to avoid undesired interferences in drinking water
treatment processes and impaired water quality. With membranes that have a
molecular weight cut-off of 1 – 5 kD (see Table 2), which means NF and the
lower range of UF, the necessary removal is achieved for colour and other
components. This is illustrated in Figure 4, showing rejection values for
various membranes and parameters in Norway [4]. It can be seen that if
removal of colour and TOC (NOM) is the prime target for the treatment, NF
is the best process. But it is not desirable to remove the scarce minerals in soft
water and therefore many NOM removing plants operating on soft water
with moderate colour (<40 ppm Pt) apply tight UF membranes. For simplicity
all NOM plants operating on such water sources are called NF plants for
simplicity, because all major properties of these plants are the same. It can
also be seen that iron is efficiently removed as this compound is primarily
bound in organics and as hydroxides, whereas manganese and partly calcium
are soluble and show low removal efficiency. It is not desirable to remove
hardness and trace minerals in surface waters in the actual areas, as the water
is naturally soft. Such solutes are beneficial with respect to health and
corrosion.
100
Fe
Rejection [%]
80
60
Colour
40
Mn
Ca
TOC
20
NF
UF
0
0.5
1
5
10
50
Nominal membrane pore size [nm]
100
Figure 4. Typical rejection of various parameters from various Norwegian sources [4].
It is evident that NF is able to remove most actual components from natural
surface waters. If the source water is seawater, brackish water or ground
water, tighter membranes are needed and that usually means RO. But NF can
and are widely used for softening and NOM may be a concern also in this
application, both as undesirable water colour and because NOM is a known
fouling material for the membranes. If softening is the treatment purpose the
use of tight NF or RO membranes is necessary. But besides the higher
operating pressures that are needed, the fundamental mechanisms for
filtration efficiency and fouling are the same. Softening, groundwater and
brackish water filtration are applications with a comprehensive history,
especially in USA. These applications are thoroughly covered in the literature
and textbooks, like [5] and [6].
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Direct NF means:
• There is no main treatment process ahead of the membrane filter.
• This is different from membrane processes with pre-coagulation and
where the membrane process is a polishing step following conventional treatment processes in front, like coagulation, precipitation etc.
• In such application both the raw water and the design and operation
of the processes upstream of the membrane plant will influence on the
membrane plant experiences. In these cases a separate investigation of
the NF plant will give an incomplete picture.
• But the general mechanisms of membrane filtration, as presented in
Chapter 2, will still be relevant.
• In direct NF the membranes alone takes care of all necessary removal
efficiency in the plant and it is the properties of the membrane that
decides the treatment efficiency.
• All application on soft surface water is called NF in this report,
although some plants strictly spoken apply tight UF membranes with
slightly larger pores than the typical NF membranes.
Since spiral wound membranes used for direct NF cannot be backwash, any
fouling of the membrane should be avoided. In principle, it should be
possible to operate the process in a way so that this is realised, see Chapter 2.
However, two main factors make this very difficult [2]. Firstly, natural water
sources contain a wide variety of particle types, from highly soluble lowmolecular solutes to macromolecules and large particles in the micron range.
There may always be some particles that from some reason settle on the
membrane surface and starts the fouling process. Secondly, the water flux of
the membrane is very important for fouling development (Chapter 2). Higher
flux increases fouling significantly, which calls for a compromise between
membrane capacity and cost and on the other side fouling. It would be
convenient to use a simple prefilter in front of the membranes that removed
the fouling agents, which will be commented later.
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2 Filtration and fouling mechanisms
2.1 Particle characterisation
Knowledge of the hydrodynamic properties of individual particles is
necessary for an analysis of particle migration during membrane filtration.
Whether the particles are visible, molecules, colloids or aggregates does not
matter. They are all examples of particles and do also have an apparent
“molecular” or particle mass. In the drinking water business, particles are
often comprehended as visible particles, which accidentally have dimensions
comparable to or larger than the wavelength of light.
For natural water published data described by Hayes et al. [7] indicate that
particles up to a molecular mass around 100 kDa are true molecules and that
few exists above 1000 kDa. It has been found that several properties of the
particles vary with the size. As shown in [7] and [8], aromatic groups are for
example most prominent in the intermediate size fractions, whereas
polysaccharides are dominant in the largest particles. Colour occurs mostly in
intermediate and larger particles. Approximately 50% of NOM is organic
carbon.
Molecular mass is often used as a size measure, although correct values are
difficult to determine because of polydispersity. This report emphasises
hydrodynamic diameter dh as a suitable measure. This is the diameter of a
sphere having the same hydrodynamic properties as the actual particle in
Brownian diffusion, as described by the Stokes-Einstein’s equation:
DB =
k ⋅T
3 ⋅π ⋅ µ ⋅ dh
(4)
Here DB is the diffusion coefficient, k is Bolzmann’s constant, T is absolute
temperature and µ is viscosity. For membrane filtration the size and shape are
important [2]. The dominant view is that the smallest particles are spheroid
and the larger ones are random coils. There are differences between soil and
lake water and between climate zones. Studies with electron microscopy and
other methods show spherical or slightly oval particles between 2 and 20 nm.
Large particles have been described as aggregates up to 30 nm and irregular
fibrous shapes up to 3 µm. Some studies report web-like structures above
50 nm. The shapes, especially for large particles, depend on concentration, pH
and the presence of other ions. The particles stretch to more linear chains in
low concentrations, low ionic strengths and neutral pH, presumably because
of less intramolecular repulsion [9].
Particle break-up is supposed to be relevant only for particles larger than
1 µm at shear rates above 1000–2000 s-1, which is higher than common in
membrane plants. Some publications indicate that NOM structures may take
days to stabilise ([25] [26]) after some change in the conditions.
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The size distribution of particles in natural water varies between water
sources. Several sources of size distribution data have been evaluated and
Figure 5 shows some typical data from Norwegian surface water. It can be
seen that most particles are between 1 and 10 nm, but some mass is also
found in larger particles. Similar curves from other sources are referred by [7],
[12] and [13]. Water with higher salinity like carbonate typically shows fewer
large particles. It is also evident that there are more large particles in bog
water and in lakes in the spring because of increased soil drainage. The
situation can be generalised by average and typical curves for the particle size
distribution, as shown in Figure 6 [2].
Relative amount of DOC
per log unit particle diameter
100
Bog, Hellerud (Ratnawera et. al., 1998)
Lake, Maridalen
-- “ -River, Sagelva (Kootatep, 1979)
Lake, Aurevann -- “ --
75
50
25
0
5 6 789
2
3
1
4 5 6 789
2
3
4 5 6 789
10
Hydrodynamic particle diameter [nm]
2
100
Figure 5. NOM particle size distribution in some Norwegian samples (personal information [2]).
Common
shapes
Spheroids
Common
compounds
Fulvic acids and
simple organics
Prolate ellipsoids
Fibres
Humic acids1
Polysaccarides
Loose webs
All NOM
TOC per log unit particle diameter,
arbitrary skale
50
20
Spring / moore and/or
less minerals
10
5
2
1
Winter and/or
more minerals
5
0.5
1
2
5
10
20
50
100
200
Hydrodynamic particle diameter [nm]
500
1000 2000
5000
Figure 6. Generalised NOM particle size distribution for rivers and lakes and its dependence on season
and water source [2].
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2.2 Filtration and fouling
In non-fouling membrane filtration particles move away from the membrane
after the flux has brought them there. This occurs by various diffusion
mechanisms where particle size and crossflow velocity is important ([19] and
Figure 7). Shear-force diffusion is caused by particle-particle collisions and
inertial lift is a result of water-particle interaction in a velocity field. Brownian
and shear-diffusion need a concentration gradient to operate. The crossflow
layer closest to the membrane surface, where there is a concentration
polarisation and a concentration gradient is often referred to as the “film”.
The inertial lift velocity can be directly subtracted from the flux.
Concentration polarisation will develop by itself in the film until equilibrium
between flux and total diffusion is obtained.
SHEAR FORCE
DIFFUSION
~ u·d s2
BROWNIAN
DIFFUSION
~ dh-1
INERTIAL LIFT
DIFFUSJON
~ u 2·d s3
Particle
Concentrate channel
CROSSFLOW
y
x
FLUX
Membrane
Figure 7. An illustration of the driving forces that act on particles in the velocity field above the
membrane surface.
A generally successful approach to fouling is based on a mass balance for
migrating particles. However, many studies rely on simplifications that are
not valid in non-fouling NOM filtration, like only considering one particle
size and one mechanism of diffusion. It is also commonly assumed in the
literature that the membrane surface is covered with an increasingly thicker
layer of fouling material downstream in the channel. This is not correct in
non-fouling operation. One of the publications that address NOM filtration in
particular is [16], but no theoretical development was given.
Thorsen [2] presented a fundamental study of how and why fouling by NOM
occurs. This was based on a mass balance for various fractions of NOM
during filtration and was done with the condition that fouling should not
occur. The analysis was based on the classical procedure described in [17] and
[18]. A method for calculation of the fouling by NOM was developed in
which the separate calculations for several hundred particle size intervals was
done with consideration of size and other properties of each sequence and of
the mutual influence between the sequences. The influence from various raw
water qualities was part of the procedure, se Figure 6. The calculated results
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had acceptable precision for the purpose to predict the performance of
technical filtration equipment.
Figure 8 shows calculated examples for two crossflow velocities in a 1 mm
open flat filter channel. The curves show the relative concentration of various
particle sizes on the membrane surface during filtration. It is evident that the
range of particle sizes that causes significant polarisation is limited and
changes with the crossflow. Lower flow gives higher concentration and
smaller particles in the most critical range for fouling, which is the range of
particle sizes that gives highest concentrations on the membrane surface.
Similar graphs for flux show that increased flux significantly increases the
surface concentrations and reduces the size of critical particles.
NOM concentration on the membrane
surface in 1:1.59 size segments [mg/l]
10
Flux = 40 l/m2h
1
0.3 m/s
0,1
1.5 m/s
0,01
5
10
50 100
500 1000
Hydrodynamic particle diameter nm
5000
Figure 8. NOM concentration at the membrane surface as a function of particle diameter and crossflow
velocity (see boxes) in an 1 mm flat channel. The total concentration is 15.5 g NOM/l for 0.3 and 5.0 g
NOM/l for 1.5 m/s. The NOM concentration in the feed is 10 mg/l.
Figure 9 gives calculated concentrations for various distances from the
entrance to the membrane channel. The most critical particle sizes for fouling
are between 0.05 and 2 µm. This points to a NOM fraction rich in
polysaccharides, which are believed to make up the skeleton of the web-like
particles in this range. The mass balance predicts that concentration
polarisation is higher downstream in the channel. This implies that the
apparent rejection of NOM, calculated from bulk concentration will be lower
in the far end. A device that disturbs the build-up of concentration
polarisation along the membrane, like the spacer in spiral wound membranes,
seems beneficial. It can also be seen that the fouling particles become smaller
downstream. This may create an additional problem as smaller particles are
more hydrophobic and may adsorb easier to the membrane ([8], [12], [19]). In
the case shown in the figure fouling will occur downstream in the channel
beyond approximately 30 mm (with a flux of 30 L/m2h). The most critical
range of particle sizes points to a fraction of NOM that contain much
polysaccharides, which are believed to make up the skeleton of the web-like
particles in this range.
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NOM concentration on the membrane
surface in 1:1.59 size segments [mg/l]
10
Crossflow = 0.5 m/s
Flux = 30 l/m 2 h
1
1000 mm
25.0 g/l
0.1
100 mm
10.6 g/l
10 mm
3.8 g/l
0.01
5
10
50
100
500
1000
5000
Hydrodynam ic particle diam eter [nm]
Figure 9. Total (boxes) and relative surface concentration of NOM with particle diameter and distance
from the inlet (data as Figure 8 unless specified).
2.3 Causes of residual fouling in practical filtration
Experiences show that in spite of good plant experiences in general, a brown
gel-like material often accumulates on the membrane. With a flux of 2530 L/m2h in spiral wound membranes the capacity can drop more than 75%
in 2000 hours [2]. Membrane filtration of NOM has been a topic in
publications worldwide, for example [20], [21] and [22]. But many experiments typically are dead-end filtration and/or application of higher fluxes
than 25 L/m2h, which leads to fouling according to [2]. In [23] and [24] it was
confirmed that most fouling was caused by the larger NOM particles.
In [25] it was shown that at high fluxes, typically 40-60 L/m2h, crossflow
velocity, Ca concentration and the flux itself have significant influence on the
flux decline during about 50 hours operation. But at moderate fluxes around
20 L/m2h the flux decline is much slower and shows low dependence on the
same parameters. Similar results were reported in [26]. These results clearly
show that there is a critical flux.
Thorsen [2] assumed that some membrane materials have a lower affinity to
NOM and are thus less prone to adsorptive NOM fouling. This is strongly
supported by [19] and is also the general impression from a series of
experimental studies by the same author during the last 20 years. Practical
experiences in full-scale drinking water plants in Norway show that the best
spiral wound membranes can be operated for weeks with an almost constant
flux up to 20 L/m2h. From model calculations with actual conditions in these
plants the critical concentration on the membrane was found to be 4 -5 g/L.
From the experiments referred to above and other experiences by the author
the critical surface concentration for fouling therefore seems to be
approximately 5 g/L. This is illustrated in Figure 10 [2], which shows
calculated concentrations of NOM at the membrane surface in spiral modules.
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C ritical long term concentrations
for m em branes with low adsorption
N O M : 0.01 g/l
34 m il spacer
10
40
30
1
20
15
0.1
10
0.05
0.1
0.15
0.2
0.25
C rossflow, [m /s]
Flux [l/m2h]
Surface NOM concentration [g/l]
100
0.3
Figure 10. Calculated surface concentrations for spiral membranes with various fluxes and crossflow
values within the recommended range (“autumn” water) from [2]. Critical flux is the value where the
concentration on the membrane is critical.
Model calculations indicated that particles in the size range of about 0.1–3 µm
are particularly critical for fouling (previous section). A structure hypothesis
for NOM particles in this critical size range predicts web-like particles [2],
which were also assumed from morphological studies [7]. A geometrical
evaluation of the actual dimensions of the particle web and void sizes shows
that the density of organic material in the particles should be about 6-7 g/L.
From the particle shapes and measurements of diffusivity and viscosity of
NOM the apparent density of the particles in this critical range can be
calculated to approximately 10 g/L. This agrees with the conception that
these particles are web-like particles and supports the conclusion on critical
particle concentrations [2].
Fouling material that has recently been deposited on the membrane may be
able to diffuse back into the bulk zone in the membrane channel, which
means that the fouling would be reversible. This has been experienced in
tests. But NOM that stays for more than a day at this concentration will
eventually develop a collapsed structure and the density of the organic
fouling material increases on the membrane. Analyses of old fouling material
and calculation of the effect of restructuring indicate an initial organic density
during fouling deposition of about 6.4 g/L, which agreed well with a
calculation from viscosity data for various NOM fractions, which gave
approximately 7.7 g/L [2].
None of the indications above are evidences alone, but seen together they
strongly suppose that fouling starts when the particles touch each other and
limit their free movement on the membrane surface. The restructuring of
NOM particles will not be able to complete within common residence times at
the membrane in non-fouling operation. The surface concentration has
moderate dependence on the actual bulk flow concentration, indicating that
the fouling problem must be similar from one natural water source to another
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[2]. Therefore an approximate value of the critical concentration on the
membrane of approximately 5 g/L seems reasonable.
2.4 The significance of the membrane type
In [10] crossflow membrane filtration experiments are described using seven
water sources and eight different membranes, featuring a wide range in
hydrophilicity, with contact angles from 13° (regenerated cellulose) to 61°
(polysulphone). It was concluded that the ratio between flux and mass
transfer coefficient for NOM (Jw/k) was the dominant parameter deciding the
flux decline. Neither NOM hydrophobicity nor membrane type was
significant in comparison during the length of their experiments (3 days). This
supports the assumption that diffusive and convective particle transport
dominates the concentration polarization.
But pilot studies with spiral membranes over several thousand operating
hours, referred in [2], show clear differences between the membrane polymer
types. Some membrane types showed more flux decline over time than
others. These results were found in several experiments. But the decline was
not apparent until after 500 operating hours. Only the membranes with the
least flux decline could be operated close to the critical values given in Figure
10 without significant fouling.
An apparently lower critical surface concentration, and lower critical flux, for
some less hydrophilic membranes can easily be explained by additional
adsorptive fouling as reported by [19]. Many studies use polysulphone
membranes, which have a tendency to adsorptive fouling. But membrane
materials exist that show almost no NOM fouling at moderate fluxes; for
example some CA membranes give little adsorptive fouling and are also
cheaper and are easier to clean than polysulphones [2]. In these experiments
CA membranes were much easier to clean completely than PS and PVDF and
PS showed the worst residual fouling in long-term operation (one year) with
occasional membrane cleaning. It therefore makes sense to avoid adsorptive
fouling in technical filtration by using suitable membrane materials.
In one publication several types of membranes were compared regarding
adsorptive fouling and pore blocking (19). Their test solution was effluent
from a SMCP pulp mill, which contains lignins, polysaccharides and low
molecular acids similar to NOM. Table 4 shows a summary of results where
the residual flux was measured after static adsorption to equilibrium in the
concentrated solution. It can be seen that regenerated cellulose (like CE
above), TFC and modified PVDF suffer very little adsorptive fouling.
Cellulose acetates and regular PVDF have moderate adsorption whereas all
types of polysulphones show severe adsorptive fouling. This is in general
agreement with a study by [27], a study of static adsorption and fouling from
lake water. Both tests showed that adsorption and fouling with polysulphone
and acrylic membranes were significant and irreversible. They also found that
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membranes from regenerated cellulose showed negligible adsorption and
fouling.
Table 4. Residual permeability after static adsorption of compounds from pulp mill effluent [19]
Membrane material
Residual permeability
Regenerated cellulose, CE
102 %
Various polyamides, TFC
102 %
Modified PVDF
95 %
Cellulose acetates, CA (tri-acetate lowest)
53 – 91 %
Regular PVDF
82 %
Polyacrylonitrile, PAN
70 %
Polyetherimide, PEI
30 %
All polysulphones, PS, PSS, PES
29 %
Another study by [28] shows more than twice as much adsorption of humic
and fulvic acids on polysulphone membranes as on acrylonitril membranes.
This agrees with the values in Table 4. They also showed than under
convective conditions adsorption reached equilibrium in the order of a few
days. This again agrees with the experiences that fouling is reversible on a
scale up to a few days. The eventual immobilisation of fouling matter
therefore seems to involve both slow adsorption and/or aggregation.
Several studies have investigated the chemical factors that may promote
adsorption, one of them being the zeta potential and it pH-dependence. But
this topic will not be detailed here.
Using high fluxes and various membrane types in experiments with NOM
fouling, would give confusing result as fouling might have been be caused by
both physical accumulation and adsorptive fouling on the membrane. Further
it is astonishing that most published experiments continue to use membranes
that are inclined to adsorptive fouling without comparing the results with
easy cleanable membranes like CA and CE (see table for abbreviations). The
generally most popular membrane materials, like PS and TFC (PA on PS)
usually have better water permeability, which should facilitate higher fluxes
than for example CA. But this capacity is impossible to realise in long-term
operation. In that case the amount of residual fouling (after cleaning) and the
“cleanability” of a membrane are most important. But of course, higher water
permeability facilitates lower operating pressure, if this important.
2.5 Chemical factors in NOM fouling
The fouling layer from filtration of natural water with NOM can be observed
as a brown-black material with noticeable thickness. The layer seems to float
on the wet and pure white, apparently clean membrane surface beneath. A
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Fouling layer [µm] and relative permeability
comprehensive study identified polyphenols and proteins in fouling material
[29]. Their analyses also showed significant amounts of inorganics (clay,
carbonate, hydroxides) embedded in the organic material. But it appeared
that the organics were somewhat concentrated closest to the membrane,
which they interpreted as adsorptive fouling. These results are in general
agreement with a study on adsorption and desorption of NOM-similar
organics with several membrane types [30]. The source water in these tests
was generally richer in inorganics and turbidity than typical Norwegian
surface water. Fouling layer thicknesses of 30 to 50 µm were seen, which is
similar to Norwegian results shown in Figure 11 [2].
100
Pilot, 2 “ spirals, ~ 25 l/m 2h
Relative permeability
Layer thickness
75
50
25
No cleaning
Experimental cleaning
0
0
1000
2000
3000
4000
5000
Operating hours
Figure 11. Development of the permeability (broken line) and the fouling layer during on-line field pilot
test using natural coloured water, colour approx. 40 ppm Pt [2].
Another study with three US water sources that were pretreated by alum
coagulation and filtration showed enrichment of polyphenols in the fouling
material. Polysaccharides, protein and amino sugars were reduced
accordingly [31]. These results indicate show that polyphenols, the more
hydrophobic part of NOM, play a part in fouling. Norwegian experiences
indicate that iron, but not calcium, may play a part in the fouling. This agrees
with [33] who tested with a TFC membrane.
In one study NOM was fractionated and the hydrophobic and hydrophilic
fractions, as well as an unfractionated NOM solution, were used in NF tests
with high fluxes (> 50 L/m2h). The most striking result is that both
hydrophilic and hydrophobic components are essential in fouling. It seems
that it is the hydrophobic compounds that cause a fouling layer, but the layer
is glued together to form a “mat” by hydrophilic compounds. A layer of only
hydrophobic compounds acts as an accumulated mass of particles that are
still free to diffuse away from the membrane if the flux is reduced [32].
Another interesting result of this study is that more NOM is recovered from
chemical cleaning after filtration of the hydrophilic solution than after
filtration of the hydrophobic solution.
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A study of the adsorption of humic and fulvic substances on different
membranes in the presence of calcium showed that the negative surface
charge of the membranes was partly neutralised in such solutions [28]. It was
also found that increasing concentration of calcium in the solution increased
the amount of NOM that was adsorbed on the membrane. The effects were
more pronounced for humic acid than for fulvic acid and for lower pH.
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3 Fouling control
3.1 General experiences
The use of membrane filtration for treatment of coloured drinking water in
Norway was proposed the first time in Norway in 1976. The first laboratory
tests were done in a PhD study by Kootatep (1979). This study was done with
plate and frame, tubular and spiral membranes. At that time cellulose acetate
membranes were dominating the membrane market and also were cheapest,
so only such membranes were tested.
The most suitable pore sizes were soon found to correspond to nanofiltration
and tight ultrafiltration membranes. It was quite obvious from Koottateps
results that a flux level of 50 to 70 L/m2h with a tubular membrane would
lead to fouling and flux decline within a few hours. But he also found that
this early fouling (before 200 hours) could be removed with a combination of
concentrate flushing about every 10 hours and chemical cleaning about every
20 hours.
These tests were regarded as promising and were followed by more
laboratory tests at SINTEF in 1980 (Thorsen, 1981). The new tests showed that
a spiral membrane could be operated with no chemical cleaning for 24 days in
succession (7.5 hours/day) at 24 L/m2h. These tests were done in a batch
apparatus with recirculation and with a 150 µm strainer in the feed line. The
main results from the long-term tests by Koottatep and Thorsen are shown in
Figure 12 In both tests the operating conditions were approximately as
recommended by the membrane producer. The colour of the raw water was
in the range 75 – 150 mg Pt/l and the turbidity was below 1 NTU, as usual
with soft Nordic surface water. With the tubular membranes the source was a
small stream, with the spirals it was diluted bog water.
The spiral membrane maintains a steady flux without cleaning, whereas the
tubular membrane does not achieve this stability in spite of 6 instances of
chemical cleaning in 170 hours. The cleaner was a typical receipt used for
membrane cleaning in dairies. The results indicated that it could be possible
to maintain a reasonable long-term flux around 20–25 L/m2h with CA spiral
membranes. An economic evaluation from this led to the assumption that
spiral cellulose acetate membranes could compete with alternative treatment
in full scale plants up to 50–100 m3/h. That assumption was experienced to be
correct later on.
No conclusion could be reached about maximum stable flux for tubular
membranes, but 50 L/m2h seemed to be too high without frequent cleaning.
As tubular systems are more expensive, they should maintain at least that
level to be competitive with the spirals. Therefore spiral membranes were
selected for further testing in pilot scale.
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80
Flux, l/m2h
Tubular CA with
cleaning (arrows)
60
40
20
Spiral CA with
no cleaning
0
0
50
100
Operating hours
150
200
Figure 12. Laboratory tests with CA membranes to find a sustainable flux [2]. The tests were done in
small units using tubular and spiral membranes respectively.
3.2 Laboratory and pilot experiments
3.2.1 Spiral wound membranes versus capillary membranes
Published studies [34] have shown a significant difference in the properties of
open membrane channels and channels with a spacer, like spiral wound
membranes. In good spiral wound membranes there is significant mixing of
the crossflow solution for each mesh in the spacer, whereas in open channels
like capillaries the concentration polarisation film grows undisturbed all the
way down the channel. This difference can be tested and in a pilot experiment
two 1-m 2”-spiral modules (about 1 m2 membrane) with identical membranes,
were operated in parallel [35]. One module had a common diamond spacer
and the other had a spacer made of corrugated plastic, featuring a flow
pattern similar to capillaries. The experiment was performed on-line in a fullscale membrane plant for surface water (Trondheim, Norway). The feed was
taken from the upstream side of the full-scale membrane modules, which
gives slightly concentrated raw water. Permeate samples were taken close to
the inlet, from the middle section and close to the outlet. The fluxes were kept
equal in each section
Examples of measured (symbols) and calculated permeate concentrations are
given in Figure 13. The curves were calculated as described in [2]. The rule for
the mass transfer in spiral wound membranes, using 2/3 of the spacer mesh
width as the typical membrane length as described in [34] was used. An
increase in permeate concentration with module length is evident for the
capillary modules as expected. But the increase was slightly less than
calculated.
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10
10
Feed water TOC
9
Feed water TOC
9
8
8
CAP
6
SPI
5
4
3
1
6
SPI
5
4
3
Equal crossflow:
Spiral: 1000 l/h
Capil .: 1000 l/h
Flux : 15 l/m 2 h
Time: ca.70 min
2
CAP
7
TOC [mg/l]
TOC [mg/l]
7
Equal energy:
Spiral: 650 l/h
Capil .: 1500 l/h
Flux : 15 l/m 2 h
Time: ca.70 min
2
1
0
0
0
0,2
0,4
0,6
0,8
1
0
0,2
0,4
0,6
0,8
Distance from inlet [m]
Figure 13. Calculated (curve) and measured permeate concentration of TOC for spiral and “capillary”
modules. Left graph: equal crossflow velocity, right graph: equal energy loss for crossflow. The
membranes were identical 20 kDa CA.
The average permeate concentration in the capillary module versus the net
spacer module agrees with the model calculations and with the mesh rule for
spiral wound membranes. The spiral shows lower permeate concentrations
and therefore less concentration polarisation. It is evident that the spiral also
shows increasing surface concentration with the distance from the inlet. The
use a fraction of the mesh width as the membrane length for calculation of
mass transfer in spiral membranes therefore is not entirely correct although
fairly accurate.
These experiments used equal crossflow (left) and equal crossflow energy at
low flux (15 L/m2h). Similar tests were done with different fluxes and sample
times and the results agree with the main result discussed above. The results
show that it is more concentration polarisation and consequently more
fouling potential for capillary membranes than for spirals at equal flux.
3.2.2 Spiral wound membranes with different prefilters
In the previous chapter it was claimed that a relatively narrow range of
particle sizes around 0.1-2 µm are especially critical for fouling. By using
prefilters of different grades in front of membrane filters the validity of this
result can be illustrated. Pilot experiments with different prefilters were
performed using three 2“ cellulose spiral wound membranes (Osmonics, Inc.)
with a cut-off of approximately 8 kDa [35]. The membrane length was 1 m
and the area was 1.6 m2 per spiral wound membrane. The spacer was a
common diamond type and the membranes were operated with typical soft
surface water directly from Lake Leirsjøen (Norway). The turbidity was about
0.4 NTU and TOC was 3 mg/l. The flux was kept at 24 L/m2h. The
membranes were operated simultaneously in parallel with identical
conditions except for the prefilter. The conditions for each membrane are
explained in Figure 14.
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1
Recovery~60%
Permeate,~38 l/t
0,86 mm “diamond”-spacer
Cartridge
filter
Concentrate
~26 l/h
Module length=1m
Coloured surface water on-line
from lake, CTOC = 2.9 - 3.0 mg/l
Recirculation,~490 l/n
Crossflow,
~ 0,18 m/s
Figure 14. The apparatus used for pilot experiment with different prefilter grades, showing some details.
To test the effects of different prefiltration, the feed to the membranes was
prefiltered with cartridges of grade 0.1, 5 and 100 µm respectively (Millipore
“Polyguard”). The cartridge cut-off is specified for mineral particles, but gives
a reasonably defined separation. The experiments were run continuously for
one month with no membrane cleaning. The membranes were rinsed with
chlorinated water 2–3 times a week to avoid biological growth (20 ppm). The
fluxes were maintained nearly constant all membranes by fine-adjusting the
feed pressure (5-8 bar, ~10 °C).
The results from the experiment are expressed as relative permeability of the
membranes versus time, as shown in Figure. 15. A relative permeability of 1.0
means no fouling. It can be seen that there is no or insignificant fouling with a
0.1 µm prefilter. With 5 and 100 µm prefilter there is a 31–37% decline, which
is not acceptable for full-scale operation. That means membrane frequent
cleaning is needed. It is also evident that the difference between 5 and 100 µm
prefiltration is small. Model calculations according to [2] show that with the
actual condition there should be fouling with 5 and 100 µm, but not with 0.1
µm prefiltration. The results in Figure 15 are convincing regarding the
efficiency of fine prefiltration.
Relative permeability
1,2
1,0
0,1 µm
0,8
5 µm
0,6
100 µm
0,4
0,2
0,0
0
200
400
600
800
Time (hours)
Figure 15. Permeability at constant flux for membranes with different prefilter grades. The feed was
typical coloured surface water.
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3.3 Fouling and rejection at different plant recovery
In order to maintain a high enough crossflow to facilitate a reasonable flux in
spiral membranes, some of the concentrate is recirculated to the feed side of
the pressure vessels. This exposes the membrane to higher concentrations of
all solutes and other particles that are rejected by the membrane. It is
important to note the membrane rejects all the particle sizes around 1 µm,
which are most critical for fouling. This must be accounted for when the
fouling potential is evaluated for a given case.
The recirculation will also influence on the difference between membrane
rejection and plant removal efficiency. This is illustrated in Figure 16. The
water samples behind the data are taken from a small full-scale plat in
operation. It can be seen that the rejection is 95% for colour at start, but it
declines to 86% after 60 000 hours, most probably because of hydrolysis of the
CA membrane. At 80% recovery however, the plant efficiency for colour
removal is 88% at start and only 70% after 60 000 hours.
It should be mentioned that a membrane with a fouling layer must be
expected to show reduced rejection because the crossflow cannot reach down
to the actual surface of the membrane. It should be remembered that the
thickness of the fouling layer may be several tens of µm. Further, the
diffusion of permeable solutes can be reduced inside the fouling material so
that the membrane is exposed to a higher concentration than in the bulk flow
of the raw water. This leads to higher transport of these solutes through the
membrane.
Treatment efficiency, E (%)
100
80
60
Colour, new plant
Colour, at 60 000 hours
TOC, at 60 000 hours
Curves calculated with 50% recirculated feed
40
20
0
20
40
60
Permeate recovery (%)
80
100
Figure 16. Plant treatment efficiency with various rejection values and degrees of
permeate recovery. Rejection is the efficiency value at 0% recovery [2].
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4 Experience in NOM removal applications
4.1 Norwegian experiences with spiralwound NF membranes
Number of membrane plants for
coloured surface water in Norway
Treatment of coloured surface water with membranes was introduced in
technical plants in Norway in 1989. Since then membrane filtration of natural
surface water for drinking water has become an increasingly popular
technology in Norway (Figure 17). A few similar plants exist in Scotland,
Ireland and elsewhere. With more than 100 plants in operation in Norway,
the need to optimisation and complete the design criteria is highly relevant.
100
80
60
40
20
0
1990
1995
2000
2005
Year
Figure 17. The number of full scale plants for the treatment of coloured surface water by NF/UF in
Norway. Plant sizes are from about 100 m3/d to about 13000 m3/d.
The concentration of natural organic matter (NOM) in the surface water
sources is particularly high in cold climates. The source of this organic matter
is mainly plant material that is slowly broken down by chemical and
microbial activity in soils and lakes. Partly decayed material is also
resynthesised and then broken down again along alternative routes. The
breakdown process is complex and slow. The resulting organic mixture in
natural waters includes a long range of chemical as solutes, colloids and
larger particles.
The main mass of particles is in the size range 1.5–10 nm (Figure 6). For
membrane filtration humic substances and polysaccharides are the most
important for fouling [2]. NOM is a potent foulant for membranes and are
also a main “pollution” of the water sources. The reasons are that they form
toxic compounds by reaction with chlorine and that they give the water a
brown colour mainly from quinones. This problem is very predominant in
Norway, which has limited groundwater resources, and therefore surface
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water is the dominant source for drinking water. But also in countries like
Canada, Ireland and Scotland the problem is significant.
The usual way to remove NOM is by chemical coagulation and sludge
separation. This is effective, but the process constantly needs optimisation
with the shifting temperatures etc. The plants need manual operation to
produce a stable water quality. Therefore some laboratory and pilot studies
were done in the 80-ies and some full-scale plants were started in 1989 – 91 in
Norway. It was concluded were CA spirals were best suited for membrane
filtration as they showed good flux stability and regained the full flux easier
than other types like TFC PA and in particular PS. A mild neutral cleaner was
for example more efficient for CA than a strong caustic cleaner for PS and this
is environmentally advantageous.
Based on very good operational experience and stable product water quality
from the first three full-scale plants, a number of new plants were started the
next ten years in northern Europe. They are all built on the same basic design
and with similar operation, as shown in Figure 18. The pretreatment in the
majority of the plants is only a self-cleaning steel cartridge with 50 µm mesh.
The turbidity of the feed water usually is below 0.5 NTU, but in some cases
like river sources with higher turbidity in some seasons, a sand filter is
installed upstream of the cartridge. A few typical examples of other source
water analyses are given in Table 5.
As the surface water in the actual areas usually is soft, there is a need for posttreatment to adjust the pH and increase alkalinity. In most small plants this is
simply done with a bed of granulated calcium carbonate with an adjusted
bypass to control pH. This will not give full alkalinity correction, but for small
plants simple operation is important. A Ca concentration around 8 mg/l and
a pH of approximately 8.5 is easily achieved. Corrosion control achieved by
this simple system is adequate in most cases.
Table 5. Some examples of soft surface water sources in Central Norway [2].
Water source
pH
Conductivity
mS/m
Colour
mg Pt/l
TOC
mg/l
Ca
mg/l
Fe
mg/l
Mn
mg/l
Stavsjøen
6.5
8.0
49
6.4
5.8
0.40
0.12
Trolla
6.8
7.6
50
5.3
4.2
0.16
0.013
Våvannet
6.2
3.9
30
2.7
2.0
0.17
0.014
Larskogvannet 5.7
5.8
79
8.4
1.5
0.34
0.016
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Chlorine
Bypass line
Cleaner
solution
Alkaline
filter
Pressure vessels, 6 m
Reservoir
Sand filter
Feed (optional)
Prefilter
Drain
Circulation pump
Figure 18. The layout of most surface water plants used for coloured water in northern Europe [2].
Figure 18 shows all unit operations involved in the plants and as can be seen
the plants are simple in principle. All plants seem to operate with
recirculation from concentrate to feed to maintain sufficient crossflow. For a
stable operation these three factors that must be controlled:
•
•
•
The flux shall never exceed the critical value for fouling, which means
15–22 L/m2h depending of the membrane type and the source water.
The membrane type is critical in order to avoid adsorptive fouling.
Therefore most membranes in use are selected CA types.
The membranes must be cleaned at proper intervals. This is usually
done daily with a diluted solution (approx. 0.5 g/L) of selected
chemicals. This procedure removes fresh fouling material before it
restructures and forms a bound fouling layer.
Rejection of colour and TOC [%]
100
Membrane: CA, 20 kD
colour
TOC
90
80
70
60
50
Br
La
Vå
yg
rs
gv
ga
ko
et
at
t
n
ne
øe
tn
la
va
ol
sj
at
lv
av
yl
Tr
H
St
ne
t
Figure 19. Retention of some parameters with pore size [2].
In addition to the daily cleaning, called “chemical rinse”, a main cleaning is
performed about once a year. The plants are simple but optimised, and are
based on the proper know-how. The selection of membrane retention is
dependent on the feed water characteristics. In some cases Mn must be
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User
reduced, and as this element is in true solution, a tighter membrane is
needed. Otherwise the need for colour removal will decide the membrane
cut-off. Both TOC and colour removal efficiency varies between the water
sources as illustrate in Figure 19. It should be observed that Figure 19 shows
the rejection values and not the plant removal efficiency, see the Chapter 3.
A significant fraction of NOM in soft water consists of long out-stretched and
negatively charged molecules. The charges repel each other and cause the
molecule to stretch. If the ionic strength of the solution increases the
repellence decreases and the molecules curl up and shrink in size. Therefore
the best choice of membrane will depend on the conductivity of the water [2].
This is illustrated in Figure 20, which shows how the relation between TOC
and conductivity influences the retention of TOC. The membrane with
molecular weight cut-off of 1 kDa is best suited.
Rejection for TOC, RTOC i %
100
1 kD
90
20 kD
80
70
Typical
”colour”membranes
60
0
0.5
1.0
1.5
2.0
2.5
TOC / conductivity, [mg/l] / [mS/m]
3.0
Figure 20. Retention as a function of membrane molecular weight cut-off and the TOC/conductivityrelation, data from [15].
A total of approximately 150 plants of this type have been started world-wide
by 2005, about 70 % of them in Norway. Figure 21 shows a photo of a
Norwegian plant with a capacity of 8000 m3/d at Frøya public water supply
that was started in 1997. Most plants have capacities in the range 100–
15 000 m3/d. Operating cost of these plants is slightly higher than
conventional chemical treatment with coagulation in spite of their simple
design, but are still often preferred because of a more stable water quality and
less manual operation.
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Figure 21. A 6000 m3/d plant for colour removal with CA spiral membranes at Frøya in Norway (photo:
T. Thorsen).
4.2 Configurations with tubular membranes
Tubular membranes are often considered to be too expensive per m2 of
membrane area to be used for water treatment. Such modules need seals for
each individual membrane tube, and each tube covers only about 0.03 m2 of
membrane area. The packing density of membrane area per plant volume is
low, causing bigger plants. But the cost of membrane filtration is closely
related to the flux. As mentioned membrane filtration of natural surface water
may suffer from severe fouling from NOM. It was actually claimed in
Chapter 3 that fouling in open channels like in capillaries and tubes are more
unfavourable than in spiral configuration as the concentration polarisation is
less disturbed than in spirals.
On some membrane types, especially CA types, the fouling material appears
as gel-like layer that floats loosely on the wet CA surface. It can easily be
wiped away mechanically. This is impossible to do with spirals, but the idea
was developed in the UK to apply the mechanical foam-ball cleaning
procedure for tubular membranes (PCI Membrane Systems). A thorough prestudy by Thames Water and PCI [36] showed that the idea was efficient. The
relatively high average flux of 24 L/m2h could be used with foam ball
cleaning at intervals of 4–6 hours. The process, called Fyne Process, operates
with a significant fouling rate, but the fouling material can be wiped away
easily from CA membranes that do not suffer from adsorptive fouling. A
sketch of the operating principle is shown in Figure 22.
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end caps with U-tubes to permit
series flow through module
Feed
Water
foam ball
foam ball catcher(s)
(hold foam ball in feed
stream ready for next clean
when flow changes direction)
Reject/Waste
72 membrane tubes in series (6 shown)
Filtered Water
Figure 22. The operating principle of foam ball cleaning [36].
The higher cost of the tubular membranes makes them most suitable for
smaller plants. But the reliability of the system now has resulted in close to 50
installations, mainly in Scotland, with plant capacities ranging from 10 to
780 m3/d [36]. The process uses PCI Membranes’ C10 module and CA202
membranes. The foam ball cleaning and minimise the use of chemicals. An
example of such plants is shown in Figure 23.
Figure 23. Typical tubular membrane plant for coloured water (courtesy of PCI Ltd).
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5 Groundwater and softening applications
Nanofiltration processes are capable of removing hardness, heavy metals,
NOM, particles and a number of other organic and inorganic substances in
one single treatment step. NF membranes have a reasonable high rejection of
bivalent ions whereas the rejection of monovalent ions is moderate to low.
Operating pressure is typically in the range of 5-30 bar. The process will be
adequate for surface and ground waters with high concentrations of total
dissolved solids (TDS), i.e. more than 500 mg/L, but with low NaCl
concentrations.
Nanofiltration membranes have properties in between RO and UF
membranes. In Table 6 the rejection of RO, loose RO, NF and UF membranes
is compared for a number of substances. The most distinctive features of
typical NF membranes are:
•
•
•
The rejection of bivalent or higher charged anions, like sulphate (SO42-)
and phosphate (PO43-) is practically total. Multivalent cations are
retained to a higher extent than monovalent cations.
The rejection of sodium chloride (NaCl) varies from about 70 % down
to 0 %.
The rejection of uncharged dissolved materials in solution depends
mostly on the size and shape of the molecule.
Table 6. Comparative rejection values for RO, loose RO, NF and UF. (Osmonics, Inc.)
Species
RO (%)
Loose RO (%)
NF (%)
UF (%)
Sodium chloride
99
70-95
0-70
0
Sodium sulphate
99
80-95
99
0
Calcium sulphate
99
80-95
0-90
0
Magnesium sulphate
>99
95-98
>99
0
Sulphuric acid
98
80-90
0-5
0
Hydrochloric acid
90
70-85
0-5
0
Fructose
>99
>99
20-99
0
Sucrose
>99
>99
>99
0
Humic acid
>99
>99
>99
30
Virus
99.99
99.99
99.99
99
Protein
99.99
99.99
99.99
99
Bacteria
99.99
99.99
99.99
99
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5.1 Hardness
Water hardness is caused by soluble ions of the alkaline earth metals, calcium,
magnesium, strontium and barium. The hardness of natural waters is mainly
formed by calcium and magnesium, since strontium and barium rarely occur
in substantial concentrations. Water hardness is not a health risk, but it is
unwanted for several reasons:
•
•
•
•
•
It reduces the effect of soap and detergents used in laundry and dish
washing. The amount of hardness minerals in water increases soap
and detergent consumption, thus adding to costs and environmental
burden.
Clothes laundered in hard water may look dingy and feel harsh and
scratchy.
Hard water may cause visible deposits on surfaces.
Heated hard water forms a scale of calcium and magnesium minerals
in boilers.
Water flow may be reduced by deposits in pipes.
Different units of measures are used to indicate the hardness of water.
• mmol/L
• mg/L CaCO3 equivalent
• German degrees (odH); one German degree corresponds to 10 mg/L
CaO
5.1.1 Types of hard water
A common distinction is made between temporary and permanent hardness.
Temporary hardness, e.g. Ca(HCO3)2, is hardness that can be removed by
boiling or by the addition of lime (calcium hydroxide). Boiling, which
promotes the formation of carbonate from bicarbonate, will precipitate
calcium carbonate out of solution, leaving the water less hard on cooling.
Hardness that cannot be removed by boiling, e.g. hardness associated with
gypsum, is called permanent hardness. Water hardness can be categorized
according to Table 7.
Table 7. Water hardness categories.
Soft
Moderately soft
Slightly hard
Moderately hard
Hard
Very hard
0-20 mg/L as calcium
20-40 mg/L as calcium
40-60 mg/L as calcium
60-80 mg/L as calcium
80-120 mg/L as calcium
>120 mg/L as calcium
Depending on pH and alkalinity, hardness above about 200 mg/L as CaCO3
(11odH) can result in scale deposition, particularly on heating. Soft waters
with a hardness of less than about 100 mg/L (6odH) have a low buffer
capacity and may be corrosive to water pipes.
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Magnesium and calcium are essential elements for the human body, but the
intake via drinking water accounts only for 5-20% [37]. A number of
epidemiological studies have shown an inverse relationship between water
hardness and cardiovascular disease in men. In most studies the calcium
concentration has shown the strongest correlation, but the magnesium
content of the water has been indicated as the most significant correlating
factor in some Canadian studies [42]. The World Health Organization (WHO)
has reviewed the evidence and concluded that the data were inadequate to
allow for a recommendation for a level of hardness [37]. There is some
indication that very soft waters may have an adverse effect on mineral
balance, but detailed studies were not available for evaluation.
Some evidence exists suggesting that drinking extremely hard water might
lead to an increased incidence of urolithiasis. The occurrence of drinking
water containing as much as 500 mg/L of calcium is, however, rare. Thus,
there appears to be no firm evidence that water hardness causes generalized
illness effects in humans.
The maximum allowable concentrations (MAC) set forward by the EU
directive are 50 mg/L for magnesium and 250 mg/L for sulphate. The guide
level (GL) figure for calcium is 100 mg/L [45].
5.1.2 Traditional softening methods
Traditional methods for water softening include ion exchange, lime softening
and pellet softening.
In ion exchange, the water is passed through an ion exchange resin. During
the passage calcium, magnesium and other bivalent or higher charged metals
are exchanged with sodium or potassium ions from the resin. When the
mediums capacity is exhausted, it is regenerated. Ion exchange leads to
increased levels of sodium or potassium in the drinking water.
Lime softening is a relatively simple process, with low to moderate capital
cost for high flow rate applications, but the hardness reduction is limited to a
minimum calcium concentration of about 20 mg/L. The process requires the
addition of large amounts of lime and acid, and produces large quantities of
sludge that requires disposal.
5.2 Scaling
A serious problem in NF systems and a limiting factor for its proper
operation is membrane scaling. Scaling or precipitation fouling occurs in a
membrane process whenever the ionic product of a sparingly soluble salt in
the concentrate stream exceeds the solubility product. Inorganic foulants
found in NF applications include carbonate, sulphate and phosphate salts of
divalent ions, metal hydroxides, sulphides and silica. The most common
constituents of scale are CaCO3, CaSO4 · 2 H2O and silica. Other potential
scalants that are rarely found are BaSO4, SrSO4, Ca(PO4)2 as well as ferric and
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aluminium hydroxides. As with other types of fouling precipitation, fouling
reduces the quality and the flux of the membrane system. The problem is
usually aggravated in attempts to increase the water recovery. Then the
increasing concentrate salt concentration may result in supersaturation, in
particular very close to the membrane surface. Scaling frequently leads to
physical damage of the membranes due to the difficulty of scale removal and
to irreversible membrane pore plugging.
In brackish and hard waters, CaCO3 and gypsum are the most common
scalants for which pre-treatment should be considered.
Calcium sulphate
The most common form of calcium sulphate scale that precipitates at room
temperature is gypsum (CaSO4 · 2 H2O). Gypsum is approximately 50 times
more soluble than CaCO3 at 30oC. The effect of temperature (in the range of
10-40oC) or pH on gypsum solubility is negligible.
One source of sulphate ion in NF applications is the addition of sulphuric
acid to the feed in order to control CaCO3 precipitation. This method of scale
control can lead to calcium (or barium and strontium) sulphate precipitation,
if excessive amounts are used for pH control. Alternatively, hydrochloric
acid, which does not contribute to scaling, may be used for pH reduction.
Calcium carbonate
The potential for CaCO3 scaling exists for almost all well, surface and
brackish water when such water is concentrated. Calcium carbonate forms a
dense, extremely adherent deposit and its precipitation in an NF plant must
be avoided. It is by far the most common scale problem.
For quantification of the tendency to precipitate calcium carbonate the
Langelier Saturation Index is frequently used.
Silica
Amorphous silica is one of the major fouling problems in NF systems. Its
solubility at room condition is 120-150 mg/L in the pH range 5-8 and it
increases significantly with pH at values higher than 9.5. Furthermore, silica
solubility increases significantly with temperature. Thus, in usual water
treatment operations silica concentration is limited to approximately 120150 mg/L, the excess precipitates as amorphous silica and silicates.
5.2.1 Scale control
It is economically preferable to prevent scale formation, even if there are
effective cleaners available for scale removal. Scale often plugs membrane
feed passages, making cleaning difficult and very time consuming. There is
also a risk that scaling will damage membrane surface. Several methods of
scale control are employed in nanofiltration:
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•
•
•
changing operational parameters
acidification
antiscalant dosage
Operational parameters that can be changed are recovery, crossflow velocity,
temperature and pressure. Reduced recovery will reduce the concentration
ratio and thereby the risk of precipitation. Increasing pressure and reduced
crossflow will enhance CaSO4-nucleation and precipitation. However, a
reduction of the recovery may be in conflict with the production target of the
facility.
Antiscalants are surface active materials that interfere with precipitation
reactions in three primary ways:
• Threshold inhibition: it is the ability of an antisclant to keep
supersaturated solutions of springly soluble salts.
• Crystal modification: it is the property of an antiscalants to distort
crystal shapes, resulting in soft non adherent scale. As a crystal begin
to form at the submicroscopic level, negative groups located on the
antiscalant molecule attack the positive charges on scale nuclei
interrupting the electronic balance necessary to propagate the crystal
growth. When treated with crystal modifiers, scale crystals appear
distorted, generally more oval in shape, and less compact.
• Dispersion: dispersancy is the ability of some antiscalants to adsorb on
crystals or colloidal particles and impart a high anionic charge, which
tends to keep the crystals separated [47]. The high anionic charge also
separates particles from fixed anionic charges present on the
membrane surface.
5.2.2 Antiscalants
Common
antiscalants
are
sodium
hexametaphosphate
(SHMP),
diethylenetriamine-penta-methyl phosphonic acid (DTPMPA) and 1hydroxyethylidene-1,1-diphosphonic acid (HEDP).
During the past two decades new generations of antiscalants have emerged
commercially, in which the active ingredients are mostly proprietary mixtures
of various molecular weight polycarboxylates, polyacrylates and
polyacrylamide. Some times a combination of at least two different types of
antiscalants are used for optimum results. Combinations of antiscalants and
acid dosing may also be effective.
Antiscalant dosages range from 2-5 mg/L. Due to the high molecular weight
and negative charge the antiscalant will have a high rejection.
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5.3 Case studies
5.3.1 Florida
In Florida, about 87% of public water is produced from groundwater [38].
Commonly, these water supplies are classified as being hard, having
relatively high concentrations of calcium. Many of the supplies also have
substantial dissolved organic colour, hydrogen sulphide and iron. Until about
1985, essentially all municipal groundwater treatment plants in Florida
practising softening used the lime softening process and, in many cases
relatively high dosages of chlorine for disinfection and bleaching-out colour.
With the introduction of more stringent drinking water standards,
particularly for disinfectants and disinfection by-products, new softening
plants have favoured membrane softening over lime softening in treating
coloured groundwater. In 1995, a total membrane softening water treatment
capacity of 350 000 m3/d was installed or under construction. Table 8 presents
raw water characteristics from three locations where lime softening and
membrane softening are used.
Table 8. Groundwater characteristics at three locations in Florida.
parentheses [38].
Permeate specifications in
Ft. Myers
Shallow wells by
river
Boynton Beach
Surficial aquifer
Plantation
Biscayne aquifer
230 (130)
295 (50)
315 (20)
200
265
285
70
50
60
TDS (mg/L)
480 (285)
380 (90)
420 (35)
Iron (mg/L)
0,3
0,2
1,7
75 (<5)
40 (<1)
65 (<5)
0,3
1,5
Trace
Raw water
source
Hardness as
CaCO3 (mg/L)
Alkalinity as
CaCO3 (mg/L)
Chloride (mg/L)
Colour
H2S (mg/L)
Figure 24 shows a typical membrane softening plant in Florida. Sulphuric
acid and, in many cases antiscalants for pH and scale control, are added to the
raw water. Cartridge filters, usually rated at 5 microns, remove particles that
may foul the membranes. The softening membranes used are typically spiral
wound NF membranes. Permeate is sent to a degasifier for carbon dioxide
(pH adjustment) and hydrogen sulphide removal. Post-treatment chemicals
are added to the degasified water. These include chemicals for disinfection
(chlorine or chloramines), pH adjustment (NaOH), and often corrosion
control (inhibitor) and fluoride.
The concentrates are disposed of differently, according to site specific
conditions and regulatory requirements. Some alternatives are injection in
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deep wells, discharge to saline surface water (e.g. ocean outfall), sanitary
sewer or blending with other water for irrigation/reuse.
Antiscalants,
pH
adjustment
Degasifier
Cartridge
Disinfection
To distribution
network
Air
Boost pumps
Membrane
filtration units
Concentrate to
disposal
Clearwell
Figure 24. Typical nanofiltration softening plant [38].
The feed pressure is usually in the range 6-9 bar. A majority of the membrane
softening plants in Florida are 2-stage plants with a recovery of 80 to 90%.
This value is a practical upper limit for a 2-stage process because of the
requirement to maintain a minimum crossflow velocity in the last membrane
element in each stage to prevent fouling. As can be seen from Table 8, more
than 90% of the hardness can be removed. At the same time a substantial part
of the colour is removed, resulting in permeate with a colour of less than 5.
A cost comparison showed that lime softening plant costs and operation and
maintenance (O&M) costs were lower than for membrane softening.
However, the relative difference in costs decreased with increasing capacity
of the facility. For a facility with a production capacity of about 50.000 m3/d
lime softening O&M costs appeared to be about 15% cheaper, whereas for a
production of 4000 m3/d there was a factor 2 in favour of lime softening. To
obtain the same quality with lime treatment as with NF, however, additional
treatment steps or chemicals are needed. If some water can be bypassed
around the membranes and blended to produce water comparable to the
finished water in the lime softening plant, the cost of membrane softening can
be even lower than for lime softening. The most important advantage of
nanofiltration is the product quality, which is superior to lime treatment
because of the additional removal of colour and turbidity.
Boca Raton
A new membrane treatment facility for groundwater for the city of Boca
Raton, Florida was designed in 2000 [29]. The plants capacity was
152 000 m3/d, which is the largest of its kind in the world. In contrast to
existing plants in the region, which were mostly 2-stage, this plant would
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have a separate third stage. The last stage could be used as concentrators
either to increase recovery or to increase output. With the first two stages
producing 85% recovery and the third stage with 50% recovery the overall
recovery would be as high as 92.5%. Instead of 15% of the feed water
remaining as concentrate, now only 7.5% would remain, allowing the city to
utilize the existing ocean outfall.
The transmembrane pressure would be 5.9 bar in stage 1 and 2, whereas the
third stage would be run with a TMP of 4.8 bar.
Most of the membrane softening plants operating in Florida lower the feed
water pH to 6.0 or lower. This is much lower than needed for just carbonate
scale control if an antiscalant is used with the acid. In fact, carbonate scale can
be controlled by antiscalants alone with no acid addition. It has been found
that certain commercially available antiscalants and dispersants increase the
rate of fouling by humic acids. However, the operators of most plants,
especially when surficial aquifer groundwater is used, have found that
membrane fouling is lower when they operate with a pH of 6.0 or less. In
accordance with results from pilot studies and to avoid handling of large
volumes of acid which would require daily delivery of truckloads of acid, it
was decided to base fouling control in the new plant on antiscalants alone. An
acid system, however, would also be designed and installed for periodic
operation at low pH a few hours per week. Additionally, acid addition might
be required for operation of the third stage.
5.3.2 Mainz, Germany
Gorenflo et al. [48] studied nanofiltration of a hard groundwater with high
content of NOM. The membrane filtration was carried out at a water
treatment plant of the public works at the city of Mainz. The groundwater
was treated conventionally by aeration, deferrization and demanganation
combined with rapid sand filtration and final chlorination. Raw water feed
for a pilot unit was taken after sand filtration. The pilot plant included a
cartridge filter of 30 micron pore size and a 2.5x40” spiral wound membrane
module with internal recirculation.
The membrane used in this case was a NF200B membrane from FilmTec
which originally was developed for the nanofiltration plant at Méry-sur-Oise,
France. The main filtration characteristics of the membrane are (i) a high
rejection of pesticides and organic matter and (ii) a high passage of calcium
[40]. The molecular weight cut-off (MWCO) given by the manufacturer is
200 Dalton and the membrane surface is reported to be slightly negatively
charged. This may be a possible reason for electrostatic repulsion of
negatively charged NOM components by the membrane [41].
The transmembrane pressure used was ∆p=5.5 bar which is relatively low.
For each recovery rate the experiment was run for 7 days. Rejection is here
defined as the observed rejection
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R=(cb-cp)/cb · 100%
where cb is the concentration in the bulk solution (concentrate) and cp is the
permeate concentration.
Table 9. Rejection at different recoveries
6%
Raw
water
2.9
6.89
0.22
Bulk
R (%)
30%
Bulk
R (%)
3.1
7.9
0.22
96.8
96.7
>90.9
4
9.2
0.36
97.1
96.3
83.3
20.4
46
1.6
98.2
98.0
98.0
87.5
88.2
69.8
108.9
67.6
223
73.3
114.7
12.3
128
101
122.2
12.4
ND
ND
78.9
88.9
81.4
15.9
ND
ND
74.9
86.7
329.6
55.3
582
506
77.5
90.3
96.7
95.3
Recovery
DOC mg/L
UVA m-1
Vis at 436 nm
m-1
Conductivity
mS/m at 25oC
Ca2+ mg/L
Mg2+ mg/L
AOX-FP µg/L
THM-FP µg/L
85%
Bulk
R (%)
It is seen from Table 9 that the rejection seems to be slightly higher at high
recovery. The DOC and UV-absorption at 254 are rejected almost completely.
Calcium rejection was higher than expected from the producer’s
specifications. This was due to the high concentration of multivalent anions
(SO42- in raw water: 122 mg/L, rejection 94.7%) and to possible complexation
of Ca2+ with humic substances. Magnesium was rejected significantly better
than calcium which is a consequence of the stronger hydration of the Mg2+ion.
The AOX (adsorbable organic halogen) and THM (trihalomethane) formation
potential were rejected by more than 95%. Even at the highest recovery (85%)
no scaling was observed. The investigated module showed no significant
fouling (flux decline less than 2%) within the 4 weeks operation period. The
fact that so little fouling was observed is probably a result of the very low
DOC-content in the water and the extended pre-treatment.
The specific flux at 85% recovery was calculated to 5.6 L/(m2 h bar) at 25oC at
an average flux of 30.8 L/(m2 h).
5.3.3 Spain
A softening plant with a capacity of 21 000 m3/d of NF permeate and a blend
capacity of up to 30 000 m3/d was installed in Bajo Almanzora, Andalusia for
the production of potable water [43]. Raw water from the Bajo Almanzora
dam is characterized by excess sulphate, calcium, magnesium and TDS. The
NF softening process aims at taking Mg2+ to less than 50 mg/L as ion and
total dissolved solids down to levels allowed by the Spanish Sanitary
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Guidelines and to lower the sulphates to less than 250 mg/L. The original
treatment facility previous to the softening plant consisted of the following
•
•
•
•
•
•
•
Pre-ozonation
Mixing chambers
Clarifiers
Sand filtration
Treated water pre-chlorination and disinfection system
Final ozonation chamber
Potable water reservoir
Figure 25 shows a general flow sheet of the plant. The configuration of each
rack is in two different arrays of 44 and 20 pressure vessels respectively. Each
PV contains six FilmTec NF membranes type NF70-345. Each train operates at
a recovery rate of 70%. Filtered water from existing pre-treatment is subject to
dosing of antiscalant, sodium metabisulfite to reduce free chlorine. Hydrogen
chloride and antiscalant is added in order to prevent precipitation of
sparingly soluble salts. Each of the three trains has a 5-micron cartridge filter
ahead of the NF membranes. The NF solution was preferred to RO due to
lower pumping costs and high reject of divalent ions. The general operational
cost savings are in the range of 10-22% compared to low pressure RO and
conventional RO respectively.
Suck-back tank
NF racks
Process
pumps
Water from
existing
pretreatment
Train 1
Cartridge filter
5 micron
1st array
44 PVs
(6el./PV)
Filtered
water
storage
Mixer
2nd array
20 PVs
Permeate
tank
Brine
disposal
to sea
Reserve
HCl
Train 2
Postozonation
chamber
Bisulfite
Train 3
Antiscalant
To regulation storage
and water distribution
Cleaning & flushing
system
Figure 25. General flowsheet of the NF70 plant, Bajo Almanzora. (After [43])
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Table 10. Key performance parameters of the NF70-345 plant
Feed water temperature, oC
Recovery, %
Feed pressure, bar
Product flow, m3/h
12
65-70
8-10.2
215-280
Table 11. Composition of feed and permeate.
Parameter
SDI *)
Feed
Permeate
4,2-4,5
TDS, mg/L
2000-2200
90-138
Conductivity
231,8-237
12,6-19,4
Ca2+
278
7,45
Mg2+
111
2,49
Na+
193,6
36,49
K+
9,4
1,99
HCO3-
82,9
8,74
SO42-
1107
21,2
Cl-
252
56,5
NO3-
4,2
2,3
SiO2
1,7
0,81
*) SDI (Silt Density Index)
5.3.4 England
At the water works at Debden Road, Saffron Walden, intermittent low levels
of pesticide in the borehole water required an appropriate treatment for its
removal [44]. A second treatment goal was a reduction of the calcium
hardness by approximately 50%. Previously this was achieved by ion
exchange softeners installed in 1947.
Nanofiltration with DOW NF200 membranes was chosen to achieve both
requirements. A cost comparison between this solution and a combined
process consisting of basic ion exchange/GAC filtration had shown that the
capital costs were about equal. However, the operation costs for
nanofiltration is lower than for the combined process ion exchange/GACfiltration. An important feature with the chosen membrane is that it only
partially removes hardness so that the only post-treatment necessary is CO2stripping and security chlorination.
The plant feed water is a non-karstic groundwater with an average calcium
hardness of 320 mg/L as CaCO3, containing intermittent low levels of
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pesticides (<0.3 µg/L). Treatment targets were to remove pesticides to below
the standard of 0.1 µg/L, and to obtain a calcium hardness of between 150
and 180 mg/L CaCO3.
The nanofiltration plant with a design capacity of 125 m3/h was in service by
the end of 1996. The membranes were arranged in three arrays with a
configuration of 14-7-4, allowing a recovery of 85% (see Figure 26). The total
membrane surface area is 5574 m2, which results in an average design flux of
22.4 L/m2h. On this site, the good raw water quality in terms of suspended
solids, turbidity and SDI allowed a configuration without any clarification
(pre-treatment) except the 5-µm cartridge filter that acts as a security barrier.
The nanofiltration membranes are fed directly by the borehole pumps. Main
results for the hardness and alkalinity reduction performance are shown in
Table 11.
Figure 26. Flow diagram at the Debden Road plant.
Table 12. Main feed water and permeate characteristics.
Feed water
Permeate
Rejection (%)
Conductivity
mS/m at 20 oC
Ca hardness
mg CaCO3/L
Alkalinity
mg CaCO3/L
62-68
40-47
30-40
290-350
150-200
40-50
270-330
140-190
35-50
pH
7.0-7.2
6.8-7.0
-
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During the first year of operation of the plant, the pesticide levels in permeate
samples have been all far below the drinking water standards (0.1 µg/L). The
pesticides found in the raw water, at intermittent, low concentrations were
atrazine (<0.16 µg/L), simazine (<0.04 µg/L) and chlorotoluron (<0.1 µg/L).
5.4 Waste disposal
The operation of a nanofiltration plant generates a concentrate as well as
different wastes originating from membrane washing and disinfection. The
issue of waste disposal must be addressed as an integral part of the design
and evaluation of the process. The method of disposal will ultimately be site
specific, depending on raw water characteristic, concentration factor and local
environmental regulations.
High recovery leads to a concentrating effect of dissolved species in the feed
water, the extent of which can be estimated from the following equation:
Cf = 1/(1-Y)
where Cf is the concentration factor of ionic species and Y is the recovery.
Usually a recovery of 80% is used for drinking water production [39]. This
implies a concentration factor of approximately 5. A high concentration factor
reduces the amounts of waste, but on the other hand, the problems with
scaling become more severe as described earlier.
The use of antiscalants contributes to the total-P content in the case of
polyphosphonates, and is considered to be a compound that promotes algae
growth. In the Netherlands, several solutions for concentrate disposal have
been considered [48]:
•
•
•
•
•
Treatment of the concentrate by rapid sand filtration or continuous
filtration before discharging.
Selection of a nanofiltration membrane with a lower rejection of
sulphate.
Selection of less contaminated groundwater wells that are used as feed
water.
Discharging the concentrate near the influent or near the effluent of a
wastewater treatment plant.
Transport over several kilometers in order to discharge the
concentrate in a larger water body.
Despite these creative solutions it is expected that concentrate disposal will
become more difficult. European legislation may result in more severe
restrictions. Therefore, there is a need for technologies that either remove
specific compounds from the concentrate before discharge or technologies
that make concentrate disposal unnecessary.
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In general, concentrate disposal can be achieved by a variety of routes
including:
•
•
•
•
•
direct or indirect disposal to receiving surface water (e.g. stream, river
or lake)
discharge to saline surface water (e.g. ocean outfall)
disposal to sewerage system (WWTP)
infiltration
reuse for irrigation
For membrane plants located near wastewater treatment facilities,
concentrate disposal to a wastewater system can be a very viable option.
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6 Summary
The description of NF application to drinking water treatment in Chapters 1
to 3 is relevant for various types of source water such as soft and hard surface
waters and ground water. All natural water sources contain some NOM and
it is usually desired to reduce the concentration of this component, commonly
expressed as TOC. Additionally, especially in polishing, other organic
components are targeted for removal too, for instance micropollutants. In
softening (ground water) application, hardness removal is a primary
treatment target. Special aspects of soft water treatment, polishing and
softening are considered in Chapters 4 and 5. For all these NF applications
NOM is considered to be a potential fouling agent.
All applications share the same filtration and fouling mechanisms. NOM is
both a treatment target and an important fouling agent. Scaling is not a
general problem in soft waters, but the mechanisms and ways of controlling
concentration polarisation that promotes scaling are the same as for NOM.
Silt fouling is not thoroughly described in this report. Silt is small inorganic
particles (clay etc.) that follow the same mechanisms of concentration
polarisation as NOM particles. The critical particle sizes are similar to NOM
particles. The main difference is that higher concentrations on the membrane
surface are tolerated before fouling occurs because the number of particles is
less than for NOM at comparable concentration because of much higher
density of the particles. Silt fouling should not be confused with scaling,
which is the precipitation of sparingly soluble inorganic salts at the
membrane surface.
Pure silt fouling may still occur at high silt concentrations, tentatively above
about 5-10 mg/L. Such concentrations are not common, at least not in typical
Scandinavian surface waters. During floods there may be such problems
some places. To protect the membranes from silt fouling, very fine prefilters
are needed. This typically means dedicated microfiltration, especially by
backflushable capillary membranes. This may easily be a more complicated
and more expensive solution than to reduce the flux until fouling does not
occur, or to select a better water source. Pure silt fouling is very difficult to
remove and the potential for this should be investigated before the
installation of a membrane plant is decided.
In some cases silt in lower concentrations may be trapped in a NOM fouling
layer on the membrane. Such combined fouling can be a problem because
membrane cleaning can disintegrate NOM fouling, but silt may remain on the
membrane. This should be controlled by good flushing after cleaning,
otherwise a similar situation as with pure silt fouling may gradually develop.
To date fouling control is the single main challenge for all the applications.
The practice in Norwegian plants shows that direct NF is applicable to soft
surface water because fouling is largely prevented by flux control and
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minimum necessary cleaning. Still problems occur in some plants. These are
suspected, from experiences so far to be of various causes:
•
•
•
•
•
Combined silt and NOM fouling as described above
Not optimal operation, like inadequate crossflow and flushing after
cleaning
NOM combined with biofouling, possibly promoted by nutrients in
cleaner remainings
Use of water sources with unusually large NOM particles from stirred
bottom sediments or direct inflow of soil drainage
Spacer clogging from bottom sediments or loosening pipe wall fouling
upstream
It has been shown that fine prefiltering, below 5 µm, will prevent fouling.
This confirmed that there is a range of particle sizes with hydrodynamic
diameters in the range 0.1–3 µm that are the main cause of (particle) fouling
(not relevant to scaling). However, prefilters that are efficient for removal of
micron-sized particles also suffer from severe performance loss due to pore
fouling and clogging by critical particles. As a consequence, microfilters need
vigorous backwashing, often combined with chemicals, to remain operative.
The net result is that such prefilters are expensive.
The role of the flux should not be underestimated. Fluxes higher than the
backdiffusion velocity will lead to accumulative fouling. In all cases there
exists a critical flux that should not be exceeded to avoid foulant
accumulation. This flux should be assessed by reliable experience or pilot
tests for more than 2 months with steady operation. There are several cases of
plant installations in which the plant supplier relied entirely on the
membrane specification in terms of permeate capacity rather than on the
critical-flux-concept.
Most membranes have a much higher theoretical capacity than possible to
realise in long-term stable operation. The application of high capacity
membranes is of no use if that capacity can not be fully utilised. This has lead
to less use of CA membranes, which have lower capacity potential, but also
low adsorptive fouling. Higher capacity membranes usually are more
expensive and CA is easier to clean in spite of lower chemical resistance, for
example too high pH in cleaners. Cleaning with the stronger cleaners
therefore is not always necessary if membranes with less adsorption of
foulants were used.
There has been an increasing focus on the proper choice of membranes with
regard to fouling tendency. This is a good thing, but there still is tendency not
to consider old fashion types like CA and other cellulose derivates. These
membranes are environmentally friendly in production and disposal,
inexpensive, have a moderate chlorine tolerance and typical life service of
50 000 hours or more. This is not the case for typical thin-film composite
(TFC) membranes with polyamide skins.
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However, in cases where tighter membranes are needed, like in softening,
polishing and the removal of manganese or low-molecular micropollutants,
the operating pressure will be higher with CA membranes. This favours
membranes with better capacity from economical reasons because less pump
energy is needed.
Fouling control is still the main challenge for NF and the most reasonable
way to control it and to maximise the plant capacity and economy, is to apply
an effective prefilter. Common strainers and microsieves have limitations
with respect to removal efficiency in the critical particle size range (0.1-3 µm).
Cheap and reliable prefilters for the 0.1–3 µm range are needed.
There is a need for better understanding the connection between source water
characterisation and proper plant design and operation, in particular how the
critical flux can be assessed. More knowledge is required to link quantifiable
raw water parameters/characteristics to membrane fouling, preferably in
form of a model.
Treatment efficiency is important in the selection of the best membrane for a
particular application. It is a question of knowing the treatment efficiency that
is needed for various parameters and to calculate the plant efficiency based
on the membrane specifications. These specifications may give rejection
rather than plant efficiency. Then the planner has to calculate plant efficiency
based on permeate recovery and the degree of recirculation. If the rejection
for special water components is not known, simple short pilot or laboratory
tests will reveal the value. This is rather straightforward, but it is essential to
have access to basic knowledge of membrane properties. There are for
example clear differences between CA and PA/TFC membranes regarding
the rejection of micropollutants of petrochemical origins, where PA/TFC
types are more efficient. More knowledge about the ability of various types of
membranes to reject specific water components is certainly useful.
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7 Conclusions
From this study the following conclusions can be drawn:
•
Nanofiltration can be used for removal of a wide range of pollutants
from groundwater and surface water in view of drinking water
production. Softening and NOM-removal are major applications, but
NF is frequently applied for the combined removal of NOM,
micropollutants, pesticides, arsenic, iron, heavy metals, sulphate,
nitrate and bacteria and viruses. Reduced THM-formation potential
can also be achieved. Full-scale installations have proven the
reliability of NF in these areas.
•
The main challenge in NF for water treatment is to control fouling of
the membrane by NOM, silt, scaling etc.
•
Regardless of other conditions there will always be a maximum flux
that can be applied in long term stable operation and therefore the flux
should be limited and not exceed this value.
•
This critical flux is almost always lower than the maximum flux
capacity of the membrane and therefore there is a significant potential
reduction in treatment costs to gain from better fouling control.
•
There is a need for better understanding of the connection between
source water characterisation and proper plant design and operation,
in particular the value of the critical flux.
•
There is a clear need for a better and cost-efficient prefilter that is
effective in the particle range 0.1 to 3 µm.
•
More knowledge of the rejection of typical and specific and important
water pollutants and groups of pollutants for various types of
membrane material would be useful.
•
For softening and groundwater applications criteria for antiscalant or
acid dosing should be developed.
•
There is a need for evaluation of waste disposal options and to assess
the environmental impact of discharge.
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