高分子概論報告
題目：membrane contractor
班級：化材三乙
學號：49940084
姓名：劉巧蓉
原理
膜接觸器是典型的晶圓廠，核苷酸與疏水性中空纖維微孔膜。 由於膜是疏水性的，並有
小氣孔，水也不會輕易通過通孔。 Thepressurerequired 迫使水進入孔隙，可以計算的 Young
－Lapace 方程式用於與疏水改性
膜（1，2）。 這種壓力通常被稱為突破壓力（等式 1）。
P =  R
應用
Boilers
Membrane Contactors have been successfully installed in hundreds of
applications worldwide and in a variety of industries such as microelectronics,
power generation,
pharmaceutical, beverage, and others. In the majority of these applications the
Membrane Contactors have been used to remove dissolved oxygen and/or
carbon dioxide for high purity water within the process facility.
Purified water
produced within a plant can be used for a wide range of applications. Use of the
water in a boiler is one such application.
In a typical boiler application, water is
supplied to a steel vessel in which a heat source is applied. The water is heated
to its boiling point producing steam. The steam is then exported to downstream
equipment. In many boiler applications the heated vessel will be fabricated
from carbon steel or similar metallurgy. It is important to minimize and control
the dissolved oxygen content in water that contacts these vessel components.
This is due to the fact that dissolved oxygen in water will react with metals in the
boiler system and cause corrosion. In order to prevent corrosion, dissolved
oxygen
is typically removed and/or controlled by mechanical means or treated with
chemicals. While chemical addition is a very effective method of reducing
dissolved oxygen, there can
be a significant cost associated with it. Chemicals have an associated cost of
procurement as well as costs required for
equipment and personnel to monitor and maintain proper chemical levels.
Storing and handling chemicals on-site also brings environmental and safety
hazards that are
prompting many facilities to reduce or eliminate chemical addition. Chemical
additives also increase the total dissolved solids (TDS) in the boiler water. Boiler
operators must monitor TDS levels and maintain them within certain
concentration
limits. If TDS levels exceed proper operating limits, scale on the boiler tubes and
other surfaces can occur. Fouling of the heat transfer surface in this manner will
affect boiler efficiency and increase operating costs of the boiler. As chemicals
are added to the boiler water, TDS levels will continue to increase.
Once TDS levels reach a pre-determined limit, the boiler must be “blown-down”.
This is essentially releasing a quantity of high TDSwater and replacing it with fresh
make-up water
until TDS levels return to proper levels.
In Figure 3 a simple flow diagram of a typical
boiler system is shown. In this drawing a
membrane contactor system is shown on the
make-up side of the process.
Figure 3: Basic Boiler System with Hybrid
Liqui-Cel Contactor System and Steam Deaerator
參考文獻
a
International Journal of Mineral Processing 96 (2010) 62–69
Contents lists available at ScienceDirect
International Journal of Mineral Processing
jou r nal h o m ep ag e : w w w. else vier. c om / l o c a t e / i j m i n p r o
Membrane contactor as a novel technique for separation of iron ions from
ilmenite leachant
E.A. Abdel-Aal a,⁎ , M.H.H. Mahmoud a, M.M.S. Sanad a, A. Criscuoli b, A. Figoli b, E. Drioli b
a
b
Central Metallurgical R & D Institute, Cairo, P.O. Box 87 Helwan, Egypt
Institute on Membrane Technology, University of Calabria, Rende, Italy
a r t i c l e
in f o
Article history:
Received 9 November 2009
Received in revised form 26 April 2010
Accepted 24 May 2010
Available online 1 June 2010
Keywords:
Membrane contactor
Solvent extraction
Iron separation
Ilmenite leaching
Titanium
a b s t r a c t
A novel system based on membrane contactor was applied for separation of iron ions during leaching of ilmenite
ore in hydrochloric acid. Separation of iron would enhance leaching of ilmenite and leads to pure titanium
products. The used membrane contactor cell consisted of two identical compartments separated by a porous flat
sheet membrane. The ilmenite leachant was separated from the residue and placed in one compartment (marked
as feed side) and an organic solution containing a selective iron extractant was placed in the other compartment
(marked as receive side). Among the several tested organic extractants, trioctylamine (TOA) was found to be
effective and selective for extraction of iron ions from solutions of a wide range of hydrochloric acid
concentrations. TOA in kerosene and 10% 1-octanole was used as a receive phase in the membrane contactor cell.
Two types of membrane materials were tested; polytetrafluoroethylene (PTFE) and polypropylene (PP) with
almost similar pore sizes of about 0.5 μm. Multi separation stages by the membrane contactor were applied by
replacing the receiving solution with a fresh one after 180 min of separation time. High Fe removal efficiency of
about 86% after 4 separation stages using 0.2 μm PP membrane was obtained. The transport mechanism of iron
was proposed mainly based on ionpair (R3NH+FeCl−4 ) formation in the aqueous–membrane interface and its
diffusion to the organic pulp through the membrane. The separation of the aqueous and organic solutions by a
membrane in this contactor technique overcame the common drawbacks of the solvent extraction such as losses
of the organic reagent, emulsion formation and delay of phase separation. In other procedures, the ilmenite slurry
suspension was placed as it is in the feed side of the cell to test the possibility of the continuous separation of iron
during leaching. The Fe removal efficiency was found to be very low (8% after 3 h) due to fouling of the membrane.
A brownish precipitate was observed on the feed side of the membrane and was growing by time leading to
slowing down of the Fe transport rate. Thin Film XRD test of membrane fouling showed that Goethite α-FeO(OH),
Feroxyhyte ō-FeO(OH) and iron oxy chloride (FeOCl) are the main constituents of the precipitate. In addition,
SEM photomicrographs showed that the precipitated particles on the membrane surface are sphere in shape with
size ranged from 1 to 2 μm.
© 2010 Elsevier B.V. All rights reserved.
1. Introduction
Titanium dioxide (TiO2) is an important intermediate in the
manufacture of paints, pigments, welding-rod coatings, ceramics,
papers, and other areas of chemical industries (Diebold, 2003).
Leaching of natural ilmenite (FeTiO3) to produce TiO2 is known to be
very hard (Afifi et al., 1994; Mackey, 1994). An intensive energy
consumption approaches such as reductive pre-treatment or smelting
of the ilmenite ore are essential stages in the current industries. In our
previous work a pronounce improvement of the leaching process was
obtained by adding iron powder as a reducing agent during leaching
of ilmenite in hydrochloric acid (Mahmoud et al., 2005). This
improvement is mostly due to the reduction of ferric ions to ferrous
* Corresponding author.
E-mail address: eabde@yahoo.com (E.A. Abdel-Aal).
ions. A comparable improvement was obtained in our recent
interesting tests by continuous removal of iron ions from the ilmenite
leaching slurry to an immiscible organic solvent under similar
conditions. This was performed by mixing the ilmenite leaching
slurry with an organic solution containing an iron extractant. This
approach will enhance ilmenite leaching, save acid consumption and
minimize iron wastes. However, difficulties such as losses of the
organic solvent in the aqueous phase, emulsion formation and delay
of phase separation can make this approach inapplicable. These
difficulties could be minimized by separation of the aqueous and
organic phases by a sheet of a suitable porous membrane. The
extraction of iron ions will take place in the aqueous–membrane
interface and the extracted species will diffuse into the organic pulp
through the membrane phase. This technique, known as membrane
contactor, would present an attractive approach for continuous
separation of iron ions from leaching slurries excluding the previous
mentioned difficulties.
0301-7516/$ – see front matter © 2010 Elsevier B.V. All rights reserved. doi:10.1016/j.minpro.2010.05.002
The concept of using a membrane to bring two phases in contact to
each other was practiced widely in diverse applications (Klaassen et
al., 2008; Kieffer et al., 2008; Fabbricino and Petta, 2007;
Phattaranawik et al., 2005; Souchon et al., 2004; Cath et al., 2005;
Mandowara and Bhattacharya, 2009). In a membrane contactor
technique the membrane separation is completely integrated with an
extraction or absorption operation in order to exploit the benefits of
both technologies fully. Membrane contactor applications that have
been developed can be found in both water and gas treatment. Several
recently developed applications of membrane contactor have been
introduced in industry, namely; selective removal of heavy metals
from a galvanic process bath, organic component recovery from a
process water in chemical industry and ammonia recovery from an off
gas stream. Expansion of this technique to metallurgical processes is
considered to be novel procedures.
Industrial membrane contactors are typically hollow-fiber modules. These contactors have three major advantages and one potential
disadvantage over conventional equipment (Noble and Stern, 2003).
The advantages are: high surface area per volume, complete loading
and no flooding (avoid solvent entrainment). The disadvantage is
slow mass-transfer (the membrane may retard the mass transfer
between the aqueous feed and the organic solvent). The success of the
membrane contactors depends on the three advantages being more
important than the disadvantage.
Iron removal from acidic leach solutions is an important and
necessary step in numerous hydrometallurgical process flow sheets
for the production of pure solutions from primary or secondary
materials (Tsai, 2009; Principe and Demopoulos, 2005; Saji and
Reddy, 2001; Zhang et al., 2010; Wang et al., 2008; Gülfen et al., 2006;
Specker and Cremer, 1959). The extraction of Fe(III) from acidic
solutions by an immiscible organic solutions was one of the earliest
systems studied in inorganic chemistry. Tri-n-butyl phosphate (TBP)
has been used by several investigators for the extraction of Fe (III)
from hydrochloric acid solutions (Haggag et al., 1977; Ishimuri et al.,
1964). Yamamura et al. (1970) described a process for the extraction
of iron from ilmenite leach liquors using methyliso-butyl ketone
(MIBK) in benzene. Extraction of iron (III) at about 1 M concentration
has been carried out from hydrochloric acid solutions using TBP, MIBK
and their mixtures (Reddy and Bhaskara Sarma, 1996). A solvent
mixture consisting of 70 vol.% TBP and 30 vol.% MIBK was found
suitable to achieve faster phase separation with limited third phase
formation. The two solvents, when used together, were also found to
exert a synergistic effect on iron extraction. The formation of anionic
iron species at high acidity and chloride ion concentrations made it
possible to extract iron by the anion exchange mechanism. Primary,
secondary and tertiary organic amines are protonated in aqueous
acidic solutions, and the protonated amines can act as a liquid anion
exchanger. After considering a number of extractants, Alamine 336, a
quaternary amine, was found to be the most satisfactory reagent for
the extraction of iron from acidic leach solutions colliery spoil
materials, requiring one to three theoretical stages (Mahi and Bailey,
1985). The amount of metal extracted was found to be independent of
the initial metal concentration in the aqueous solution but very
dependent on the initial acid concentration. Meng et al. (1996) have
studied the kinetic of iron (III) extraction with primary amine and
TBP.
In this paper, separation of iron ions from acidic chloride solutions
containing dissolved iron and titanium was investigated utilizing
different organic extractants. The suitable extractant was applied in a
membrane contactor cell consisted of two compartments separated
by a porous membrane. The ilmenite leachant containing mainly iron
and titanium in hydrochloric acid was placed in one compartment
(marked as feed side) and an organic solution containing the
extractant in a diluent and a modifier was placed in the other
compartment (marked as receive side). The transport efficiency of
iron ions from the feed to the receive sides was estimated and
parameters those affecting the transport process were evaluated. The
purified solution with reduced iron content can be then recycled to
the leaching slurry to enhance the process and produce purer
titanium oxide. This separation system was also examined for
separation of iron from ilmenite slurry during leaching.
2. Experimental
Materials
Ilmenite ore
A representative sample of Abu Ghalaga ilmenite ore was provided
by El-Nasr Mining Company, Egypt. The sample ranges in size from
1 mm to 30 mm. It was crushed and thoroughly mixed by ring and
cone method, quartered several times and a sample of about 1 kg was
ground to 100%–200 mesh (− 75 µm) and used for the present
investigation.
Leachant, organic reagents and diluent
Pure hydrochloric acid (35%) was supplied by El-Naser Chemical
Company, Egypt and used as a leachant for ilmenite ore. The used
solvent extraction reagents are Tributyl phosphate, TBP (Henkel),
Trioctyl phosphine oxide, TOPO, (Fluka), Di-2-ethyhexyl phosphoric
acid, D-2-EHPA (Luoyang Zhongda Chemical Co., Ltd.), Methyl
isobutyl ketone, MIBK (Qingdao Lasheng Corporation Ltd.), Trioctylamine, TOA (Henkel), and Aliquat 336, Aliq (Aldrich). Basic
information about these reagents are shown in Table 1. Kerosene
(El-Naser Chemical Company, Egypt) was used to dilute the active
organic extractants. Titanium IV stock solution was prepared by
dissolving a known weight of pure potassium titanyl oxalate and 5 g
of ammonium sulfate in 10 M HCl and diluted to 1 l with distilled
water. Iron (III) stock solution was prepared by dissolving a known
weight of pure ammonium iron (III) sulfate in 10 ml of concentrated
HCl and diluted to 1 l with distilled water.
Equipment and apparatus
Chemical analysis of titanium and iron in ilmenite ore, leachant and
wash liquors was performed according to the standard methods using
CECIL 7200 Double Beam UV Spectrophotometer and 3100 Perkin
Elmer Atomic Absorption Spectrometer. Leaching process was carried
out on multi stage bases. Leaching was carried out using a 500 ml three
necked glass reactor provided with a thermometer and a reflux
condenser. The reaction slurry was agitated with a mechanical stirrer.
The reactor was immersed in a thermostatically controlled water bath.
The used filtration system was a Büchner-Type porcelain funnel
connected with a vacuum pump and a suction gauge. The filter medium
used was polypropolyene (pp) filter cloth with 75 μm (200 mesh,
ASTM Standard) opening. The used membrane contactor system
(schematically demonstrated in Fig. 1) is consisted of two identical
120 cm3 tanks connected with pp or polytetrafluoroethylene (PTFE) flat
sheet membrane (pore sizes: 0.45 and 0.2 µm, diameter: 3 cm, porosity:
about 75%). The PTFE membrane (Whatman) has polypropylene grid
as the support material. The pp membranes were provided by The
Institute on Membrane Technology, University of Calabria, Rende,
Italy.
Procedure
Acid leaching and filtration
Acid leaching of ilmenite ore consists of two main stages, namely
digestion (leaching) and solid/liquid separation (filtration). In the
digestion stage: acid leaching of ilmenite ore was carried out using a
500 cm3 three necked glass reactor provided with a reflux glass
condenser and a mechanical agitator with a Teflon coated stirring rod.
A 140 ml 20% hydrochloric acid was placed in the reactor and a 20 g
ilmenite ore was added. The reactor was heated to 70 °C using a
E.A. Abdel-Aal et al. / International Journal of Mineral Processing 96 (2010) 62–69
64
Table 1
Basic information of organic reagents used in the study of iron extrication.
Extarctant
Commercial name
Formula
MW
Tributyl phospate
TBP
C12H27O4P
266.31
Trioctyl phosphine oxide
TOPO
C24H51OP
386.63
Di-2-ethylhexyl phosporic acid
D-2-EHPA
C16H35O4P
322.42
Methyisobutyl ketone
MIBK
C6H12O
100.16
Trioctyl amine
TOA
C24H51N
353.67
Trioctylmethylammonium chloride
Aliquat 336, TOMAC
C25H54NC
446.25
thermostatically controlled water bath. A steering speed of 400 rpm was
applied to keep the slurry suspended during the leaching experiment.
After 1 h of the reaction period, the slurry was filtered off. The leachant
(filtrate) was then purified by the membrane contactor technique to
separate the iron contents. When required, a second stage acid leaching
of the ilmenite residue at comparable conditions was carried out using
the purified leachant. The slurry was filtered and washed 3 times with
3% HCl solution to separate any residual leachat from solids. The filtrate
and wash liquor were analyzed for total Fe content.
Structure
Solvent extraction
In solvent extraction experiments, the 10% diluted extractant in
kerosene was placed in a 250 ml cylindrical glass vessel and mixed
with an equal volume of the aqueous phase of synthetic solution
containing iron and titanium ions in a thermostat shaker (GFL Model
1083) for the period of time required. After phase separation a sample
from aqueous phase was withdrawn and used for chemical analysis of
metal ions with Atomic Absorption Spectrometer.
Membrane separation
Membrane separation tests were carried out using a membrane
contactor cell (as schematically shown in Fig. 1) under a controlled
operating conditions. In the feed side, a 100 ml of ilmenite leachant was
placed. The leachant contained Fe ions (4000 ppm) and excess of Ti ions in
about 16–18% HCl. In the receive side, a 100 ml 30% TOA, 10% octanol in
kerosene was placed. The used membranes were PP or PTFE. The duration
of each membrane contactor experiment was 3 h and the temperature was
fixed between 50 and 60 °C. For chemical analysis of Fe, samples of 1 ml
were taken from the feed solution after 0.5, 1, 1.5, 2 and 3 h of membrane
separation time. Each sample was mixed with 1 ml concen- trated HCl and
the mixture was completed to 500 ml with distilled water. The removal
efficiency of iron was calculated as follows:
Fe removal efficiency; % = ðFeÞA −ðFeÞB = ðFeÞA
Fig. 1. A schematic diagram of the used membrane contactor cell.
× 100
where: (Fe)A is the total weight of Fe in the used filtrate solution
before separation, and (Fe)B is the total weight of Fe in the same
solution after separation.
E.A. Abdel-Aal et al. / International Journal of Mineral Processing 96 (2010) 62–69
65
Speciation diagram
The speciation diagram of iron (III) in acidic chloride solutions was
constructed using Stabcal software developed by Dr. Hsin H. Huang,
Department of Metallurgical and Materials Engineering, Montana Tech
of the University of Montana, Butte, Montana.
3. Results and discussions
Speciation of iron(III) in hydrochloric acid solutions
It was important to define the form of Fe (III) ions that exist at
conditions similar to those examined in this study. The speciation
diagram of iron (III) in acidic chloride solutions was constructed using
Stabcal program and results are presented in Fig. 2. It can be seen that
+ are predominant
cationic iron(III) species such as Fe3+, FeCl2+, FeCl
2
at low HCl concentrations and the neutral species FeCl3 is the main
one at moderate to high HCl concentrations. Moreover, the anionic
−
species FeCl4 is formed at moderate HCl concentration and became
the main one at concentrated solutions. It is believed that the levels of
both acid and chloride ion concentrations are the reasons for the
stability or instability of any species of iron at such conditions. The
variation of the form of iron(III) with concentration of HCl can suggest
the corresponding variation of the extracted species. The form of the
predominant species at a specific HCl concentration can thus predict
the extraction and transport mechanism.
Selection of the suitable extractant of Fe from hydrochloric acid
It is important to study the solvent extraction behaviour of iron
ions to choose the suitable extractant and extraction conditions before
the application in the membrane contactor system. Several organic
reagents can extract iron ions from hydrochloric acid solutions where
the extent of extraction depends mainly on the nature of the organic
extractant, chloride ion and acid concentrations. Organic reagents
belong to different chemical families; namely organophosphorus
compounds, ketones and amines, were tested for extraction of iron
ions from synthetic solutions containing 1000 ppm each Fe (III) and Ti
(IV), at a wide range of hydrochloric acid concentrations and the
results are plotted in Fig. 3.
Extraction with organophosphorus compounds
The investigated reagents were neutral organophosphorus compounds: tributyl phosphate (TBP), and trioctylphosphine oxide
(TOPO), and an acidic organophosphorus compound: Di-2-ethyhexyl
phosphoric acid (D-2-EHPA). TBP and TOPO are commercially
available and used widely as industrial extractants, especially for
reprocessing of nuclear fuels. The extraction of iron with TBP was low
at HCl concentration beyond 10% and sharply increased with
increasing acid concentration (Fig. 3). Almost quantitative extraction
was obtained at 20% HCl. TOPO showed a much better extraction of
Fig. 3. Effect of HCl concentration on extraction of iron using different extractants.
iron compared with that of TBP at HCl concentration less than 20%.
Only 10% HCl is sufficient to achieve almost complete extraction of
iron with TOPO. The three octyl groups in TOPO cause better
hydrophobic characteristics than the three butyl groups in TBP. This
may be the reason for lesser solubility of TOPO in the aqueous phase
and hence better iron extraction.
The extraction of iron by the nneutral organophosphorus reagents
(such as TBP and TOPO) is known to follow the salvation mechanism
where undissociated, electrically neutral molecules are formed and
extracted to the organic phase. According to Specker and Cremer
(1959) the extracted species with TBP from 4 M HCl is FeCl3.3TBP that
is created through adduct formation of TBP and FeCl3, and from 6 to
9 M HCl is HFeCl4.2TBP that is created by ion association of H+, FeCl4−,
and TBP. Ion association with cationic chloro complexes of iron is also
+
−
proposed by Pospiech et al. (2005) where FeCl2 Cl TBP is created by
association of FeCl2+, Cl− and TBP. These forms are consistent with the
predominant iron(III) species at the mentioned levels of
HCl, as shown in Fig. 2.
On the other hand, the extraction of iron ions by acidic
organophosphorus compounds follows the cationic exchange mechanism where an ionpair is formed and extracted to the organic phase.
The cationic exchange process involves the exchange of metal cations
with the hydrogen ions of the reagent dissolved in the organic phase.
Of the acidic organophosphorus compounds, D-2-EHPA is the most
widely used extractant (Gupta, 2003). The extraction of iron with D2-EHPA reached about 60% at 1% HCl and dropped at more
concentrated acid solutions (Fig. 3) due to the lower dissociation of
the acidic organic extractant. The extraction is negligible at more than
10% HCl. Therefore, the tested acidic organophosphorus extractant is
not suitable for iron extraction from concentrated acid solutions that
are typically used during ilmenite leaching. Moreover, the stripping of
iron (III) from the loaded D-2-EHPA is known to be hard and special
techniques such as reductive stripping can be used to solve the
problem (Lupi and Pilone, 2000).
Extraction with ketones
The investigated extractant is methyl isobutyl ketone, MIBK, which
has wide applications in metals extraction through the solvation
mechanism. As shown in Fig. 3, the extraction of iron by MIBK was in
low levels along with the studied range of HCl concentrations. The
extracted species to the organic phase in this case appeared to be
HFeCl4 (Reddy and Bhaskara Sarma, 1996). MIBK suffers from higher
solubility in the aqueous phase compared with the other common
solvent extraction reagents. Other drawbacks such as low flash point
and high vapor pressure restricted its industrial applications.
Fig. 2. Speciation diagram of Fe(III) in hydrochloric acid.
Extraction with amines
The investigated amines were trioctylamine, TOA, (a tertiary
amine) and \Aliquat 336, Aliq, (a quaternary ammonium salt). The
E.A. Abdel-Aal et al. / International Journal of Mineral Processing 96 (2010) 62–69
extraction of iron(III) with TOA and Aliq was increasing with HCl
concentration, reached about 97% at 10% HCl and remained at this
high level at more concentrated HCl solutions, Fig. 3. Trioctylamine
can be easily protonated when contacted with acidic solutions and
then can act as anion exchanger where only the anionic species can be
extracted from the aqueous solutions by exchange of chloride ion
with the metal anionic species. Aliq has a permanent anion exchange
characteristic which does not depend on the acidity of the aqueous
phase. The lower extraction of iron(III) with the two extractants at
dilute acid concentrations may be due to the scarcity of the anionic
iron(III) species. Iron (III) forms cationic, neutral, and anionic chloro
complexes based on the acidity and chloride ion concentrations as
shown in Fig. 2. Mahi and Bailey (1985) suggested the formation of an
extracted species with 1.3 mol of Alamine 336 per mole of iron(III)
based on equilibrium diagrams. The extraction data are explained in
terms of the formation of both polymeric and polynuclear iron
species. At concentrated HCl solutions, TOA and Aliq can extract iron
effectively and selectively out of any other positively charged and
neutral existing species such as those of titanium. This feature gave
these amine extractants the advantage of selectivity over other tested
extractants. The equilibrium between the iron (III) species will be
shifted towards the formation of the extractable anionic species until
all iron is separated from the aqueous solution to the organic phase.
The two reagents, TBP and TOPO, can extract titanium ions
together with iron ions from acidic conditions and then they are not
considered as a selective extractants of iron. Trioctylamine is
commercially available and was found to be a suitable extractant of
iron (III) because it can extract anionic chloro species of iron ions
selectively. Thus, TOA was used to test the selective transport of iron
ions out of ilmenite leachant in a membrane contactor system.
Possible mechanism
membrane contactor cell
of
Fe
transport
through
the
The solvent extraction of iron ions from ilmenite leach solution
was found to encounter several difficulties such as phase separation,
emulsion formation and solvent losses. When a porous membrane is
placed between the aqueous and the organic solutions, the extraction
of iron ions will be achieved without the above mentioned difficulties.
This will be carried out in a membrane contactor cell using TOA in
kerosene and 10% 1-octanol as shown in Fig. 1. The 1-octanol was used
as a modifier to eliminate third phase formation. The possible
mechanism that can take place during the transport through the
membrane contactor system, and reactions occur at aqueous–liquid
membrane interface, is schematically presented in Fig. 4A and B and
can be defined as follows:
1. Amine protonation. In the first stage, the TOA will be protonated
when contacted with HCl in the feed side–membrane interface
forming the protonated amine R3NH+Cl−.
R3 N
−
þ
m
þ HCla →R3 NH Clm
ð1Þ
where the subscript “a” denotes the aqueous solution in the feed
side and “m” denotes the membrane phase.
2. Diffusion of the protonated amine. The protonated TOA will be
diffused in the organic phase in the receive side until the whole
amount of TOA is protonated.
þ
—
diffusion
þ
−
R3 NH Clm → R3 NH Clorg
ð2Þ
where the subscript “org” denotes the organic solution in the
receive side. This solution will then act as a liquid anion exchanger.
−
3. Ionpair formation. The anionic chloro iron species, FeCl
,
4
will
interact with the R3NH+Cl− where the Cl− will be replaced with
the FeCl4 forming the ionpair R3NH FeCl4 . The latter will be
−
+
−
66
Fig. 4. Scheme of iron transport through the membrane contactor cell.
extracted to the membrane phase where the Cl− will be back
transported from the membrane phase to the aqueous phase in the
receive side.
þ
−
−
R3 NH Clm þ FeCl4 →R3 NH
−
þ
−
FeCl4 m þ Cla
ð3Þ
4. Diffusion of the formed ionpair. The formed ionpair will be more
soluble in the organic phase than in the aqueous phase. This will
obviously lead to the transfer of the neutral species from the aqueous
to the organic phase. It is then diffused to the organic pulp in the
receive side.
þ
—
diffusion
þ
−
R3 NH FeCl4 m → R3 NH FeCl4 org
ð4Þ
The sequence presented in stages 3 and 4 will be continued
until
−
all active sites in the amine are fully occupied with FeCl4 or all
iron is transported from the feed to the receive side.
Effect of membrane type on Fe removal efficiency
The membrane type must be carefully chosen to retard the mass
transfer as little as possible. Polypropylene (PP) membrane is a low
cost, resistant under extreme pH conditions, hydrophobic and
insoluble in most solvents. Polytetrafluoroethylene (PTFE) membrane
is more costly, more hydrophobic, extremely inert and suitable for
processing of aggressive streams (Nunes and Peinemann, 2006). Two
experiments were carried out with the ilmenite leachant in the feed
side using membranes made of PTFE and PP of micron pore size 0.5
and 0.45, respectively. A solution consisting of 30% TOA in kerosene
and 10% octanol was used as a receive solution. The Fe removal
efficiency was determined and plotted against time in Fig. 5. Iron
removal efficiency was fast increasing with time during the first 2 h
and then slowed down at longer time. In general, low iron removal
efficiencies of about 8% and 12% were achieved with PP and PTFE
membranes after 3 h, respectively. The transport profile showed that the
extent of iron transported was always slightly higher with PTFE
membrane than the PP membrane.
The low removal efficiency of both membranes may be attributed
to their large pore size. This may cause losses of the extracted iron in
the aqueous feed side which in return decreases the value of the
removal efficiency of iron. The slight lower transport behaviour of iron
E.A. Abdel-Aal et al. / International Journal of Mineral Processing 96 (2010) 62–69
67
Fig. 5. Effect of membrane type on cumulative Fe removal efficiency at different times.
with the PTFE membrane may be attributed to the lower masstransfer of iron ions due to the more hydrophobicity of this kind of
membrane. The PP membrane was used in the following experiments
since there is no big difference in iron transport between the two
tested membranes.
Effect of membrane pore size on Fe removal efficiency
Two experiments were carried out using the ilmenite leachant in
the feed side of the membrane contactor cell using PP membrane of
0.45 and 0.20 μm pore size, separately. The results are given in Fig. 6.
Iron removal efficiency was much higher (40%) with PP membrane of
0.2 μm pore size than that of 0.45 μm pore size (8%) after 3 h. This
improvement may be attributed to little losses of the iron extracted
species in the aqueous phase side.
Effect of number of membrane separation stages on Fe removal
efficiency
A series of experiments was carried out using multi membrane
separation stages. The used membrane was PP with 0.2 μm size, the
original ilmenite leachant was used in the feed side and 30% TOA in
kerosene and 1-octanol was used in the receive side of the membrane
contactor cell. After 3 h of the first membrane separation experiment
the loaded organic solution was replaced with a fresh organic solution
of the same original composition and the produced partially purified
leachant was used as it is in the second membrane separation
experiment. This procedure was repeated in the third separation
experiment. The results of the cumulative efficiency of transported
iron are given in Fig. 7. Iron removal efficiency was very fast during
the first membrane separation stage and slowed down with
increasing the number of stages. Also, iron removal efficiency was
relatively faster in the first 2 h than at longer separation time. This
Fig. 7. Effect of membrane separation stage numbers and times on cumulative Fe
removal efficiency.
may be attributed to the decreasing levels of Fe content in starting
solution. Table 2 shows the Fe contents in the feed solutions in weight
percentage at the different transport times. With decreasing the Fe
content in starting solution, the Fe removal efficiency was decreased.
Cumulative iron removal efficiency of about 78% was achieved after
only 2 h at the third stage of membrane separation process. The Fe
content was reduced from 0.42% to 0.09% and to 0.06% after 3 h in the
third and fourth stages of membrane separation, respectively. Almost
86% of Fe could be separated from ilmenite leachant after the fourth
stage of transport. It is clear that the highest Fe removal rate was
obtained in the first stage of membrane separation after 30 min. On the
other hand, Ti losses are very small (b 0.1%) after the fourth stage as
there is a very limited percentage of Ti diffused to the organic solution
in the receive phase. By this way, most of iron in the ilmenite leachant
could be separated without any significant separation of Ti. Also, the
separation of the aqueous and organic solutions by a membrane in this
contactor technique overcame the common drawbacks of the solvent
extraction such as losses of the organic reagent, emulsion formation
and delay of phase separation.
Effect of suspended solids on membrane separation
The transport separation of iron ions was tested using two aqueous
media in the feed side; the first is the clear aqueous phase after
filtration of slurry of ilmenite leaching (ilmenite leachant) and the
second is the leaching mixture with the accompanied suspended
solids (ilmenite slurry). These experiments were performed to study
the effect of the presence and absence of suspended solids on Fe
removal efficiency. The used membrane was PP with 0.2 μm size and
30% TOA in kerosene and 1-octanol was used in the receive side of the
membrane contactor cell. The results of Fe removal from the clear
solution and slurry are given in Fig. 8. It is obvious that, Fe removal
efficiency using ilmenite leaching slurry was much lower (about 8%)
than using the clear solution (about 40%) after 3 h of one stage
membrane separation.
Table 2
Fe content of membrane separation stages.
Fig. 6. Effect of membrane pore size on cumulative Fe removal efficiency at different
times.
Membrane
separation
time, min
Fe content,%
Stage 1
Stage 2
Stage 3
Stage 4
0
30
60
90
120
180
0.42
0.32
0.29
0.27
0.25
0.25
0.25
0.20
0.18
0.16
0.14
0.14
0.140
0.120
0.117
0.111
0.090
0.090
0.090
0.086
0.080
0.073
0.064
0.064
E.A. Abdel-Aal et al. / International Journal of Mineral Processing 96 (2010) 62–69
brownish at longer times. This developed precipitate was found not
dissolved in kerosene, TOA, ethanol, octanol and acetone but mostly
dissolved in acids such as HCl. The transport of iron according to the
proposed mechanism that is described in Section 3.3 depends on
amine protonation and its diffusion and ionpair formation and its
diffusion. The formation of such dense precipitate on the feed sidesurface of the membrane suggests the slow diffusion of the extracted
species in the membrane and organic pulp. This may lead to
accumulation of the extracted species in the layer close to membrane
surface in the feed side causing saturation followed by precipitation.
More faster shaking speeds, more thinner membrane sheets and more
porosity of the membrane material could solve this problem.
Fig. 8. Effect of membrane separation time on cumulative Fe removal efficiency from
solution and slurry.
Fig. 9. X-Ray Diffraction analysis of membrane precipitate.
A brownish precipitated material was found in the feed side of the
membrane surface after the transport experiment when the ilmenite
slurry was used. Thus, the noticeable decrease in the removal
efficiency of iron in the presence of the suspended solids may be
explained based on the extensive fouling of the membrane surface.
The fouling material clogged the pores of the membrane and
prevented the contact between the aqueous and the organic phases
which in turn adversely affected the transport of iron. This fouling
precipitate was slight and yellowish in colour in the first 20 min of the
membrane separation time and became more dense and
darker
Characterization of the membrane precipitate
The precipitated material formed on the surface of membrane of
the above experiment was separated and characterized using
Scanning Electron Microscope (SEM) and X-Ray Diffraction spectroscopy (XRD). The obtained results are given in Figs. 9 and 10. Fig. 9
shows Thin Film XRD chart of the membrane precipitate. It can be
seen that the precipitate composed mainly of Goethite α-FeO(OH),
Feroxyhyte ō-FeO(OH) and iron oxy chloride FeOCl where no any
phase contain Ti or ilmenite. Fig. 10 shows SEM photomicrographs of
precipitated solids on membrane surface. It is clear that the
precipitated particles are sphere in shape with size ranged from 1 to 2
μm. Analysis results of the precipitated particles by Energydispersive X-ray Spectroscopy (EDX) are given in Table 3. The
precipitate contains only Fe, Cl, C, and O. The precipitate material
does not contain Ti, where the Fe content represents about 60% of the
precipitate. The absence of Ti in EDX and ilmenite in the XRD analyses
indicates that the suspended solids of the ilmenite ore did not
contribute to the formation of membrane fouling. Presence of carbon
(about 12%) in the membrane precipitate may be originated mainly
from the membrane composition and possible contamination with
the organic solution. Iron oxy chloride represents about 50% of the
total precipitate.
Enhanced ilmenite leaching
After ilmenite leaching, the slurry was filtered and the filter cake
was washed with 3% HCl. The iron and titanium contents were
determined in the filtrate and in the wash liquor. The total contents of Ti
and Fe in the filtrates were 230 and 4200 ppm, respectively. This
solution was purified using the membrane contactor cell. The used
membrane was PP with 0.2 μm size, the ilmenite leachant was used in
the feed side and 30% TOA in kerosene and 1-octanol was used in the
Fig. 10. SEM photomicrographs of precipitated solids on membrane surface.
68
E.A. Abdel-Aal et al. / International Journal of Mineral Processing 96 (2010) 62–69
69
Table 3
Analysis results of membrane precipitate Energy-dispersive X-ray Spectroscopy (EDX).
Element
%
Fe
Cl
C
O
60.2
17.7
12.6
9.5
receive side of the cell. The leaching process was repeated using the purified aqueous solution after 4 stages of membrane separation and using the
residue from the first digestion. Again, after the second ilmenite leaching, the slurry was filtered and the filter cake was washed with 3% HCl.
The total Ti and Fe contents in the filtrate after the second leaching stage were 850 and 4000 ppm, respectively. This showed that excessive amounts
of titanium and iron can be dissolved from the ilmenite ore in the second leaching using a purified leachant after carrying out separation of large part
(about 85%) of iron content by the membrane contactor. Thus, the separation of iron would enhance leaching of ilmenite and leads to pure titanium
products.
4. Concluding remarks
1. Trioctyl phosphine oxide (TOPO), trioctylamine (TOA) and Aliquat
336 (Aliq) gave high Fe Extraction efficiency with 10% HCl concentration (N 97%). In addition, Tributyl phosphate (TBP) gave high Fe Extraction
efficiency (N 99%) with 20% HCl concentration. On the other hand, Di-2-ethyhexyl phosphoric acid (D-2-EHPA) gave considerable extraction
from low acid concentrations and the extraction dropped at higher levels of HCl concentrations. Methyisobutyl ketone (MIBK) gave low Fe
extraction efficiency at all studied HCl concentrations.
2. Trioctylamine (TOA), and Aliquat 336 (Aliq) in kerosene gave almost quantitative Fe extraction in more concentrated HCl (16– 18%).
These extractants are more advantageous because they are selective for Fe anionic species out of other neutral and cationic ones present in solution.
TOA in kerosene and 10% octanol was used as a suitable organic phase for continuous transport of iron from ilmenite leachant in the membrane
contactor cell.
3. The proposed transport mechanism includes amine protonation, diffusion of the protonated amine, ionpair formation, and diffusion of the
formed ionpair.
4. About 78% of iron could be transported through the membrane after 3 stages separation and this was increased to about 85% after the 4th
separation stage. The used membrane was polypropylene (PP) with 0.2 μm pore size.
5. Using ilmenite slurry in the feed side gave very low Fe removal efficiency (8.1%) due to fouling of membrane (coating with solid particles and
preventing further separation). The slow diffusion of the extracted species may cause saturation in the layers close to the membrane sides
followed by precipitation.
6. Thin Film XRD of membrane precipitate showed that Goethite α- FeO(OH), Feroxyhyte ō-FeO(OH) and iron oxy chloride (FeOCl) are the main
constituents.
7. SEM photomicrographs of the precipitated solids on membrane surface showed that the precipitated particles are sphere in shape with size
ranged from 1 to 2 μm.
8. EDX results of the precipitated particles indicated that Fe, Cl and O contents are 60.2%, 17.7% and 9.5%, respectively. Iron oxy chloride represents
about 50% of the total precipitate.
9. Ilmenite leaching was enhanced using the purified leachant after removal of about 85% of contained iron by the membrane contactor system.
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B
Journal of Membrane Science 401–402 (2012) 175–189
Contents lists available at SciVerse ScienceDirect
Journal of Membrane Science
jo u rn al
hom epa ge: w w w . e l s e v i e r . c o m / l o c at e / m e m s c i
Mathematical modeling and cascade design of hollow fiber membrane contactor
for CO2 absorption by monoethanolamine
Somnuk Boributh a , Wichitpan Rongwong a , Suttichai Assabumrungrat b , Navadol Laosiripojana c ,
Ratana Jiraratananon a,∗
a
Department of Chemical Engineering, King Mongkut’s University of Technology Thonburi, Bangkok 10140, Thailand
Department of Chemical Engineering, Faculty of Engineering, Chulalongkorn University, Bangkok 10330, Thailand
c
The Joint Graduate School of Energy and Environment, King Mongkut’s University of Technology Thonburi, Bangkok 10140, Thailand
b
a r t i c l e
i n f o
Article history:
Received 4 October 2011
Received in revised form 6 January 2012
Accepted 31 January 2012
Available online 9 February 2012
Keywords:
Carbon dioxide
Cascade design
Membrane contactor
Membrane wetting
a b s t r a c t
The absorption of CO 2 from the gas mixture (CO 2 –CH4 ) by polyvinylidenefluoride (PVDF) hollow fiber
membrane contactor using monoethanolamine (MEA) as the absorbent was performed. The mathematical model has been developed to predict the absorption performance. The model is validated with the
experimental results for estimating the wetting ratio (x*) as the function of liquid velocity and MEA concentration. The suitable hollow fiber membrane module with effective fiber length of 50 cm is selected for
the design of multistage membrane contactors. The absorption flux of multistage membrane contac- tor is
simulated based on the value of x* obtained from the experiments. The three-stage cascade design is
selected to compare the system performance with different gas and liquid flow patterns. The results of the
simulation show that the individual gas flow (G-ID) gives higher performance compared to the gas flow
in series (G-IS) for all operating conditions studied. The three different flow patterns of liquid including (i)
liquid flow in series (L-IS), (ii) liquid flow in series with splitting (L-ISS) and (iii) liquid flow in series with
recycle (L-ISR) are compared. At low MEA concentration (0.25 M), the L-ISR can improve the system
performance at low liquid velocities, while L-ISS shows the highest performance at high liquid velocities.
For the system with high MEA concentration (1.0 M), L-ISR can improve the performance at low to
moderate liquid velocities, whereas L-ISS does not improve the system performance at any liquid velocity.
© 2012 Elsevier B.V. All rights reserved.
1. Introduction
Natural gas is the clean energy that has been widely used for
several purposes such as transportation and electricity generation.
The main constituents of natural gas are CH4 and CO2 with traces of
other impurities. The removal of CO2 from natural gas prior to use is
necessary since it reduces the heating value, causes pipe line corrosion and takes up volume in the transportation. Generally, CO2 can
be captured by amine based absorption in various mass transfer
contactors such as packed, plate and spray columns. Nevertheless, these conventional contactors have limitations in operation,
i.e., flooding, foaming and channeling. In addition, high capital
and operation costs are also the major disadvantages of these
contactors for the real application. To overcome these problems,
the hybrid process called gas–liquid membrane contacting process
has been developed. This process provides the additional advantages in terms of its high contact area per unit volume, modularity
∗
Corresponding author. Tel.: +66 2470 922; fax: +66 2428 3534.
E-mail address: ratana.jir@kmutt.ac.th (R. Jiraratananon).
0376-7388/$ – see front matter © 2012 Elsevier B.V. All rights reserved.
doi:10.1016/j.memsci.2012.01.048
and compactness [1]. Therefore, the research on CO2 absorption
using membrane contactors has been intensively studied by many
researchers [2–6].
For gas–liquid membrane contactors, the microporous
hydrophobic hollow fiber membranes are generally employed as
the phase barrier. The membrane provides the contact between
gas and liquid phases without dispersing one phase into another
phase. However, the membrane adds the additional resistance to
the overall mass transfer, especially in case of the partially
wetted mode where the membrane pores are partially pene- trated
by the liquid absorbent. In the operation, the non-wetted mode is
preferred to achieve highest absorption performance. Although
the highly hydrophobic membrane is used, the par- tial wetting
can occur owing to many factors. The degree of membrane
wetting depends
on
the
structural
characteristics of
membrane, operating pressure of gas and liquid phases and
the nature of liquid absorbent in contact with the mem- brane
surface. Additionally, the direct contact of the polymeric
membrane surface with the liquid absorbent over a prolonged
operation time can lead to morphological change and gradual
membrane wetting [7,8]. Many researchers have addressed the
negative effects of membrane wetting on the process performance
[9–11].
The mathematical models have been widely proposed to investigate the degree of membrane wetting and the effect of membrane
wetting on the absorption efficiency of the membrane contactors.
Wang et al. [12] developed the two dimension model to predict
CO2 absorption and validated the model with the experimental
results. They reported that the reduction of the overall mass transfer coefficient could reach 20% when the membrane pores were
5% wetted. Mavroudi et al. [9] proposed the first-order expression
to describe the membrane resistance change with time of physical
absorption of pure CO2. Khaisri et al. [10] developed the mathematical model for investigating mass and heat transports based on
resistance-in-series model to describe the influence of wetting on
the process performance. Lu et al. [11] presented the model taking
into account the effect of membrane pore size distribution and pressure drop of liquid flow in hollow fibers on the membrane wetting.
These models employed similar basic equations and correlations
for describing the mass transfer. However, they presented different
focuses and the experimental data used to validate the models were
also obtained from different experimental conditions, including gas
compositions and liquid absorbents. Although it may be possible to
apply the results from some mathematical models for designing the
membrane gas absorption process, from our knowledge, there has
been no much report in the literature. Recently, we published the
suitable design of the membrane contacting process for physical
absorption of CO2 [13]. The main objective of the present work is to
propose a simple mathematical model to validate the experimental
data in order to obtain the important parameter (wetting ratio, x*).
From a separation point of views, a membrane contactor is one of
the unit processes, which should be designed as a multistage cascade module in the real application. The study of multistage cascade
membrane contacting process could bring about important practical criteria of the process design for improving the overall efficiency
of the process.
In this work, the absorption of CO2 from the gas mixture containing CH4 using PVDF hollow fiber membrane contactor by MEA was
performed. The effects of operating conditions including gas/liquid
flow rates, MEA concentration and gas composition on absorption
flux were studied. The mathematical model has been developed to
predict the system performance and the membrane wetting. The
wetting ratio (x*) is estimated by validating the model with the
experimental results. The suitable module length is employed in
the multistage cascade design of membrane contactor based on the
parameter (x*) obtained from the experiments. The performance of
the
with
different
gas recycle
and liquid
is compared.
Thesystems
effects of
splitting
(˛) and
ratioflow
(ϕ )patterns
are investigated.
The
results of process design at different operating conditions are also
presented.
Fig. 1. Mass transfer regions and resistance-in-series in partially wetted mode for
gas–liquid membrane contacting process.
gas-filled pores, and membrane for liquid-filled pores, respectively.
H represents Henry’s constant, and do, di, dln, and d∗l are the outer,
n
inner, and logarithmic mean diameters of non-wetted
and wetted
membranes, respectively. E is enhancement factor accounting for
the effect of chemical reaction on absorption determined by the
following:
(Ha∗)2
−
+
E = 2(E∗
1)
∞ −
4(E∗
∞
1)
−
2
E∗ (Ha∗)2
+∞∗
1
(E
1)
+
∞ −
(2)
where Ha* and
E∗ are Hatta number and asymptotic
∞ infinite
enhancement factor, respectively. These parameters can be estimated by the followings:
k2DA,L CR,L
Ha∗ =
(3)
Lk
where k2 is the second-order reaction rate constant (between CO2
and MEA), Di,L is diffusivity of species i in liquid solution and CR,L is
the concentration of absorbent.
E∗
2. Theory
(Ha∗)2
∞
=
1+
1
Di,L
CR0,L DR,L
v C
D
R
i,L
3
D
(4)
R,L
i,
L
Mass transfer in gas–liquid membrane contactor
The operation of gas–liquid membrane contactors can be classified into 3 modes including non-wetted (dry) mode, partially
wetted mode, and wetted mode. Generally, the operations of membrane contactors are performed in a partially wetted mode in which
there are four mass transfer regions as shown in Fig. 1 and the
resistance-in-series model for partially wetted hollow fiber module
can be written following Eq. (1):
1 = 1 +
1
KL∗dln
EkL di
Ek∗
d
M
+
1 +
1
HkM dln
HkG d0
where R is a stoichiometric coefficient of reaction, DR,L is diffusivity
of absorbent in liquid phase, CR0,L and Ci,L is the concentration of
absorbent at inlet and concentration of species i in liquid phase
which is in equilibrium with gas phase, respectively.
In addition, Eq. (1) can be expressed in term of resistances as
shown in Eq. (5);
Rtot
(5)
=
M
RL
+
R∗
+
RM
+
RG
(1)
ln
where KL is the overall mass transfer coefficient and kL , kG , kM , and
k∗M are the mass transfer coefficients of liquid, gas, membrane for
where Rtot, RL , R’M, RM and RG are total resistance, liquid phase
resistance, wetted membrane resistance, non-wetted membrane
resistance, and gas phase resistance, respectively.
177
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Individual mass transfer coefficients
3. Mathematical model
Gas side mass transfer
For the gas flow in the shell side, Yang and Cussler [14] proposed
the correlation to predict the gas side mass transfer coefficient (k G)
as the following;
kG dh
Sh =
dh
= 1.25
D
i,G
0.3
3
0.93
L
Re
Sc
(6)
where Sh is Sherwood number, Sc is Schmidt number, Re is Reynolds
number, Di,G represents the diffusivity of component i in the gas mixture
(see Appendix A for the calculation), L is the fiber length and dh is
the hydraulic diameter.
Liquid side mass transfer
For liquid flow in the lumen side, Yang and Cussler [14] also
proposed the liquid side mass transfer coefficient (kL).
Sh =
kL
di
di Re Sc
L
=
D
i,L
1.62
0.33
(7)
Eq. (7) is valid for Gz > 20.
Membrane mass transfer
For the non-wetted zone of the membrane pores, the membrane mass transfer coefficient can be estimated by the following
equation [15];
DG,eff εM
kM =
Model development
Mathematical model has been developed to predict the chemical absorption of CO2 by gas–liquid membrane contacting process
using MEA as the absorbent. The model is validated with the
experimental results for estimating the value of wetting ratio (x*).
Additionally, the absorption performances of multistage membrane contactor with different cascade designs are also predicted
by the model simulation based on the wetting ratio as a function of
operating conditions obtained from the experiments. The model is
derived based on the following assumptions;
1. Steady state and isothermal conditions
2. Constant total pressure in gas phase (no pressure drop) along the
fiber length
3. Ideal gas behavior
4. No axial mixing in phases
5. Membrane with uniform properties including pore size, tortuosity, porosity, thickness, hydrophobicity and pore
size distribution
6. No property change of membrane over a period of operation time
7. The wetting ratio (x*) depends only on liquid velocity and MEA
concentration.
The schematic diagram describing the mass transfer of CO2 from
gas phase to liquid phase through the hydrophobic microporous
membrane is depicted in Fig. 2. For a small element with fiber length
Ai , the mass balance of CO2 transfer from
gas z and contact area
(8)
M ıdry
phase to liquid phase is written as the following;
where ıdry is the dry thickness of the membrane, M is membrane
tortuosity and εM is membrane porosity. The gas effective diffusivity (DG,eff) is determined by the interactions between the molecules
(molecular diffusion) as well as the interactions of the molecules
with the pore wall (Knudsen diffusion). The gas effective diffusivity
is estimated using the following equation;
AL,i
=
1
1
+
DG,M
(9)
FA,i+1
−
JA,i
Ai
where FA and fA are the molar flow rate of CO2 in gas phase and
liquid phase, respectively. JA,i is absorption flux of CO2 determined
by the following;
JA,i
1
DG.eff
=
FA,i
(13)
=
KL,i
A,i
(C∗
−C )
(14)
where KL,i is local overall mass transfer coefficient, CAL,i is the bulk
concentration of CO2 in liquid phase and C* A,i is the CO2 concentra-
DG,Kn
where DG,M and DG,Kn are the Fickian’s molecular diffusion coefficient and the Knudsen diffusion coefficient of gas, respectively.
The molecular diffusion coefficient of gas (DG,M) is calculated
from the kinetic gas theory;
tion in liquid phase in equilibrium with the bulk concentration of
CO2 in gas phase given by the following equation;
PAG,i
∗
(15)
C
=
A,i
H
DG,M = 0.001858T
(10)
where PAG,i is the partial pressure of CO2 in gas phase estimated as the
following;
3/2
PM1/2
A
2
AB
˝D
where MA is the gas molecular weight, P and T are the gas pressure
and temperature, respectively. AB is characteristic length, the collision integral
D is dimensionless function of temperature which
is calculated by empirical equations (see Appendix A). The Knudsen diffusion coefficient (Dg,Kn) can be calculated by the following
equation;
=
PAG,i
yA,i
Ptot
(16)
where yA,i and Ptot are mole fraction of CO2 in gas phase and the
total pressure of gas phase, respectively. The mole fraction of CO2
can be estimated from the molar flow rate as;
FA,i
yA,i =
(17)
F + FI
A,i
,i
DG,Kn = 4850dpore
T
M
A
(11)
For the wetted part of the membrane pores, the membrane mass
where FI,i is molar flow rate of inert gas not absorbed by the liquid
transfer coefficient can be expressed as [15];
Di,L εM
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
absorbent. Therefore, FI is constant along the fiber length and is
equal to the molar flow rate at inlet feed gas (FI0).
The contact area (
Ai ) for each small element is determined
=
(12)
k∗M
ı
M wetted
where ıwetted is the wetted thickness of the membrane.
from the following:
Ai = ˝(di + x ∗ıM ) z
178
(18)
where x* is wetting ratio representing the degree of membrane
wetting.
179
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Fig. 2.
The schematic diagram showing mass balance of a small element in gas–liquid membrane contacting process.
Table 1
Membrane and membrane module specifications.
Properties
Membrane
Fiber i.d.
Fiber o.d.
m
Average pore radius
m
Porosity
Tortuositya
Module
Number of fiber
Shell diameter
Effective length
a
from deionized water and monoethanolamine (MEA, 99.8 mol%)
obtained from QrëC, Malaysia.
Descriptions
Experimental module
Designed module
800
1166
m
m
800
1166
0.08
m
0.08
4.2.
0.7
3
0.7
3
100
1.6 cm
22 cm
518
3.64 cm
50 cm
Reported by Khaisri [10].
Substitute Eqs. (14), (15) and (18) into Eq. (13) gives Eq. (19):
FA,i = FA,i+1 − ˝KL,i (di +
yA,i Ptot
z
(19)
−
H
M )
CAL,i
The average absorption flux can be determined by Eq. (20) as following:
(FA,0 − FA,N )
JA,av
(20)
=
x ∗ı
3.2.
Experimental procedures
m
Ai
Numerical solutions
For the counter flow operation, the inlet conditions of gas and
liquid phases are known, while the outlet conditions of gas and liquid are unknown. To solve this problem, the shooting method is
applied. The procedure begins by assuming the outlet concentration of CO2 in liquid phase. The simulation starts from the gas feed
inlet to the outlet. Along the axial direction, the fiber is divided into
many small elements with identical length ( z). In each element,
the set of equations (Eqs. (1), (19) and (20)) including overall mass
transfer coefficient (KL,i) and CO2 absorption flux are employed. The
outlet liquid compositions are determined by applying the mass
balance equation along the fiber length. MATLAB was employed in
solving.
4. Experimental
Materials
The polyvinylidenefluoride (PVDF) hollow fiber membranes
were purchased from Altrateck (China). The specifications of the
membrane and membrane module are shown in Table 1. Carbon
dioxide (CO2, 99.8 vol.%), methane (CH4, 99.9 vol.%) were obtained
The experimental set up of gas–liquid membrane contactor is
illustrated in Fig. 3. In this work, all experiments were performed
at a room temperature (25 ◦ C) and atmospheric pressure. The flow
rates of feed gas consisting of CO2 and CH4 supplied from the compressed gas cylinders were individually adjusted and controlled by
the mass flow controllers, (Brooks Model 4800 series). The set point
controller (Brooks Model 0254) was used to control and monitor
the flow rate of each gas entering the module. In the experiments,
the gas mixture was fed through the shell side of the membrane
module, counter-currently to the absorbent flow into the tube side
of the fibers. A peristaltic pump (I/P digital Masterflex model 759245) delivered the absorbent from the absorbent reservoir through a
rotameter to the membrane module. The inlet and outlet gas volumetric flow rates were measured by a digital gas flow meter (Bios
International, DC-lite). The inlet and outlet concentrations of CO2
and CH4 were analyzed by the Gas analyzer (Cubic, Gasboard-3200).
Before entering the gas analyzer, the moisture in the outlet gas was
removed by the water trap. Since CH4 is the flammable gas, the
methane-rich gas was vented outside the laboratory building by
the pipe line in which the flame arrestor was installed at the end.
In addition, CH4 detector (BW technologies, GasAlert MicroClip)
with response range of 0–5 vol.% was also installed to detect any
leakages and ensure safety operation.
All of the data were collected after the experiments reached the
steady state. The results of each run were averaged from four times
of sampling. The absorption fluxes were calculated from the measurements of CO2 concentration and the gas flow rate at the inlet
and outlet. The Absorption flux can be determined by the following
equation:
QG,inCAG,in − QG,out CAG,out
(21)
JA =
A
where QG,in and QG,out are the gas flow rates at inlet and outlet,
respectively. CAG,in and CAG,out are the concentrations of CO2 at inlet
and outlet, respectively. A is mass transfer area. In addition, the
removal efficiency defined as % removal was also reported. It is
estimated by the following equation:
%
=
CAG,in — QG,out CAG,out
removal
QG,in
× 100%
(22)
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
from Thai Industrial Gases PLC. The absorbent used was prepared
QG,inCAG,in
180
181
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Fig. 3.
Experimental set up of gas–liquid membrane contacting unit.
5. Results and discussions
Experimental results
Fig. 4(i) shows of the effect of liquid velocity and MEA concentration on the absorption flux. It was found that the absorption
flux increased moderately with liquid velocity, but the increase
was more significant with increasing MEA concentration. From liquid velocity 0.1–0.4 m s −1, the absorption flux increased roughly 21
and 25%, for MEA concentrations of 0.5 M and 1.0 M, respectively.
The increase in liquid velocity could increase the liquid phase mass
transfer coefficient (Eq. (7)) resulting in the increase of overall mass
transfer coefficient (KL). In addition, the driving force of the system
which is the concentration difference of CO2 at gas–liquid interface
(ii)
2.5
1.2
0.25 M
1.5
0.125 M
1
0.5
0
0.05
0.15
0.25
0.35
Liquid velocity, vl (m/s)
0.45
1
2
0.5 M
0.55
0.8
-3
2
mol/m .s)
1.0 M
CO 2 flux (x10
CO2 flux (x10-3 mol/m2.s)
(i)
and bulk liquid is enhanced with increasing liquid velocity, while
increasing MEA concentration resulted in enhanced chemical reaction rate between CO2 and MEA. The similar results were reported
in the literature [3,5,13].
The effects of gas velocity and gas composition on absorption
flux are depicted in Fig. 4(ii). It was found that increasing in gas
velocity did not affect the absorption flux as observed from the constant absorption flux. Atchariyawut at el. [16] reported the same
trend of results for absorption of CO2 from CO2–CH4 gas mixture
using PVDF hollow fiber, the CO2 flux remained constant when the
gas velocities were varied from 1.0 to 8.0 m s−1. Khaisri et al. [5]
also found that for physical absorption of CO2 using PTFE membrane, the CO2 flux remained constant while the gas velocity was
increased. Meanwhile, the absorption flux increased approximately
0.6
0.4
40 %vol. CO2
50 %vol. CO2
0.2
0
0.06
0.07
0.08
0.09
0.1
0.11
Gas velocity, v g (m/s)
Fig. 4. (i) Effect of liquid velocity on absorption flux of CO2 at various MEA concentrations (yCO2 = 0.4, QG = 0.5 l/min) (ii) effect of gas velocity on absorption flux of CO2 at
different CO2 concentrations in gas phase (CMEA = 0.125 M, vl = 0.3 m s−1 ).
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
14.3% when CO2 concentration in gas phase was increased from
40 to 50 vol.% as a result of the enhanced driving force for mass
transfer.
Membrane wetting estimation
The data obtained from the experiments were validated with the
proposed model for estimating the membrane wetting. The degree
of membrane wetting is expressed in term of wetting ratio (x*),
defined as the ratio of the length of the liquid filled pores to the
membrane thickness (x* = ıwetted/ıM). The values of x* were determined by adjusting the value (x*) in the model until the absorption
flux obtained from the simulation was very close to or equal to
the experimental data. The difference allowed between these two
values did not exceed 0.1%. In Fig. 5(i), the dots represent the wetting ratio obtained from validation of the experiments with the
model, while the trend lines are used to obtain the wetting ratio
correlation as a function of the operating conditions. It can be seen
that x* increases with liquid velocity and MEA concentration. The
increase in MEA concentration leads to significant reduced surface
tension and contact angle between the membrane surface and solution. These could increase the membrane wetting. Lu et al. [11]
reported that the membrane wetting observed from flux decline
for absorption with 1.0 M MEA was higher than that of 0.5 M MEA.
They discussed that the increase in aqueous organic solution concentration would affect the pore wetness on account of alternation
of solution properties. It is also observed that at low MEA concentration (0.125 and 0.25 M), the increase of x* with liquid velocity
is more significant compared to that of high concentrations. Due
to the distribution of membrane pore sizes, the large pores may be
penetrated by liquid absorbent [11,17]. With high MEA concentrations, the liquid absorbent can easier penetrate the large pores
compared to the low MEA concentration. However, there are limited numbers of large pores which can be wetted. Therefore, the
wetting ratio is limited to the typical value. It can be observed from
Fig. 5(i) that the values of x* for all MEA concentrations become
close at high liquid velocity. Boributh et al. [17] studied the effect
of pore size distribution on membrane wetting. It was found that at
any liquid–gas pressure difference, the membrane wetting would
occur in the pores with size equal to and larger than the criti- cal
radius. Fig. 5(ii) reveals that gas velocity and gas composition do
not affect the membrane wetting in range of the experiments
(gas–liquid flow rate ratios (QG/QL) in range of 0.45–2.0). The similar result was reported by El-Naas et al. [18]. They found that at
low QG/QL range (approximately 1.0–2.0), the membrane wetting
was quite constant. While, at high QG/QL range (2.3–3.5), the wetting decreased with increasing QG /QL . This is especially true when
the membrane has significant large pores. Therefore, the effects of
gas phase velocity and composition on wetting ratio were ruled out
and the correlation expressing x* as a function of liquid velocity and
MEA concentration was developed:
x∗
2
= (0.0725C
MEA − 0.1456CMEA + 0.0803)ln vL +
0.1062C−0.1787
MEA
(23)
In the simulation, if the value of x* estimated from above equation is
negative (at very low liquid velocity and low MEA concentration),
x* is adjusted to be zero. The above correlation is assumed to be
able to predict the value of x* for higher liquid velocity, but is in the
range of laminar flow. The system performance of the multistage
cascade membrane contacting process (presented later in Section
5.5) is simulated based on the membrane wetting obtained from this
correlation.
181
Resistance analysis
In this work, the experiments were performed at the liquid
velocities in the range of 0.1–0.4 m s−1. For the process design, it is
assumed that the wetting ratio (x*) correlation given by Eq. (23) can
be used to predict the wetting ratio with the liquid velocities up to
1.0 m s−1, which is still in the laminar flow region. By using Eq.
(23), at MEA concentration of 1.0 M, the values of x* for liquid
velocities of 0.6, 0.8 and 1.0 m s−1 are 0.1025, 0.1045 and 0.1062,
respectively, or the increase is not significant. The total resistance of
the partially wetted membrane can be determined by Eq. (1) or (5).
The gas phase mass transfer coefficient (kG) is calculated from Eq.
(6), the liquid phase mass transfer coefficient (kL) is determined
from Eq. (7), the non-wetted (kM) and wetted (k∗M ) membrane
mass transfer coefficients are determined from Eqs. (8) and (12),
respectively. The enhancement factor (E) accounting the chemical
reaction expressed in the terms of kL and k∗ M is estimated by Eq. (2).
The resistance contributions of gas, non-wetted membrane, wet- ted
membrane and liquid phases are defined as RG/Rtot, RM/Rtot,
RM∗ /Rtot and RL/Rtot, respectively and are shown in Fig. 6. For MEA
concentratio of 0.25 M (see Fig. 6(i)), the gas phase resistance is
n
around
3% and the non-wetted membrane resistance is less than 1%
for all liquid velocities. At low liquid velocities (0.1–0.2 m s −1), the
liquid phase resistance dominates the overall resistance (80–60%),
while the estimated resistances for wetted membrane are in a
range of 16–36%. For moderate liquid velocities (0.3–0.4 m s −1), the
wetted membrane resistances become comparable to the liquid
phase resistances and the wetted membrane resistance contributions are around 46–57%. At high liquid velocities (0.6–1.0 m s −1),
the overall resistance is controlled by the wetted membrane resistance (61–69%). It can be observed that the contribution of liquid
phase resistance greatly decreases with liquid velocity. In contrast,
the wetted membrane resistance significantly increases with liquid
velocity due to the increase in x*. Fig. 6(ii) shows the contribution
of individual resistance at MEA concentration 1.0 M. The gas phase
and non-wetted membrane phase resistances are about 6.5 and
1.0%, respectively, for all liquid velocities. For low to moderate liquid velocities, the resistances of wetted membrane are comparable
to liquid phase resistances. At high velocities, the wetted membrane resistance dominates mass transfer of system. By comparing
the liquid phase resistance of the system with MEA concentrations
of 0.25 and 1.0 M, it is found that the resistance contribution of
liquid phase of 0.25 M is clearly higher than that of 1.0 M. It is
because higher MEA concentration could significantly improve the
rate of reaction which is expressed in term of enhancement factor
(E) leading to reduced liquid phase resistances [19].
Suitable module length
Previous experiments on CO2 absorption by membrane contacting process have been carried out using the hollow fiber modules
with effective lengths of 10–30 cm [5,11,20,21]. These laboratory
modules provide enough mass transfer area for observing the effect
of operating conditions on the absorption performance. However,
these modules are not suitable for the pilot or industrial scales since
they provide too low contact areas to achieve high removal efficiency. Therefore, in the design of membrane contacting process for
the real application, the suitable membrane module specifications
(especially module length) should be carefully considered.
The hollow fiber modules have received wide attention due to
providing very high mass transfer area. Although, most hollow fiber
modules are designed for pressure-driven membrane processes
(rather than concentration-driven processes) such as MF and UF
processes, they can also be employed as membrane contactors.
One of the most well-known transverse flow hollow fiber modules
is the Liqui-Cel® Exra-Flow modules commercialized by CELGARD
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S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
(i)
(ii)
0.12
0.08
Wetting ratio (x* )
Wetting ratio (x* )
0.1
0.08
0.06
1.00 M
0.50 M
0.04
0.25 M
0.02
0
0.05
0.09
0.25
0.35
0.06
40 %vol CO2
0.05
0.125 M
0.15
0.07
0.45
Liquid velocity, v l (m/s)
0.04
0.06
50 %vol. CO2
0.07
0.08
0.09
0.1
0.11
Gas velocity, v g (m/s)
Fig. 5. Simulation results of (i) effect of liquid velocity and MEA concentration on wetting ratio (yCO2 = 0.4, QG = 0.5 l/min), (ii) effect of gas velocity and CO2 concentration in gas
phase on wetting ratio (CMEA = 0.125 M, vl = 0.3 m s−1 ).
Fig. 6.
Contribution of individual resistance to overall resistance (i) MEA concentration = 0.25 M, (ii) MEA concentration = 1.0 M, (QG = 0.5 l/min, yCO2 = 0.40).
LLC (Charlotte, USA). These module are 20.3–118.6 cm length. Pall
Corporation provides the modules of hydrophobic hollow fiber
membranes (PVDF, PP and PTFE) called Microza module. The available commercial module lengths are in range of 13–222.7 cm.
Gabelman and Hwang [1] recommended the commercially available parallel flow hollow fiber modules provided by various sources
which would be suitable for membrane contactors and the module
lengths are 17.8–304.9 cm.
From the above information, the module with an effective fiber
length of 50 cm is selected for the design of multistage membrane
contacting process. The selected length is also in the same range
as the membrane module used by Yeon et al. [22] for absorption
of CO2 from flue gas using pilot-scale membrane contactor (PVDF
membrane) with module length of 52 cm. To maintain the dynamic
behavior of gas flow in the shell side, the packing density of fiber
in the design module is fixed to be the same as the experimen- tal
module. The specifications of designed module are summarized in
Table 1. It is assumed that the wetting behavior occurring in this
module can be predicted by Eq. (23) because the membrane
specifications are the same as those of in the experimental module.
Design of multistage cascade membrane contacting process
The design approach is schematically shown in Fig. 7. From
the
experimental
and
modeling
results,
the
correlation for
determining the wetting ratio is obtained (Eq. (23)). The suitable of
module length and membrane characteristics used in the designed
are as shown in Table 1. The important parameters affecting the
performance of the multistage cascade membrane contactor are
investigated, i.e., number of module, MEA consumption, gas flow
pattern and liquid flow pattern. The simulation was performed for
the Reynolds number 89–890 and 4–14.23, for the liquid side and
gas side, respectively. The system performance is evaluated based
on the total feed gas which is the indicator that the system is able
to handle high feed gas capacity.
Number of modules
The module arrangement of the multistage cascade membrane
contacting process is shown in Fig. 8. The liquid absorbent is fed in
the lumen of the hollow fibers, counter currently to the gas flow in
the shell side. Fig. 9 presents the total feed gas for achieving 90%
CO2 removal for the system using different number of mod- ules.
It is found that the total feed gas increases with number of module
because the contact area is increased leading to improved
absorption capacity.
For the proposed model, the effect of pressure drop on membrane wetting as well as system performance is not included. From
the previous results, the membrane wetting is only a function of liquid velocity and MEA concentration. Therefore, to lower the effect
of pressure drop which is proportional to the module length on the
182
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Experiments
Wetting ratio (x*)
Suitable membrane module
Modeling
Multistage cascade design
Number of module
Gas flow pattern
Liquid flow pattern
Suitable process design
Fig. 7.
The design approach for multistage cascade gas–liquid membrane contacting process.
Fig. 8.
The multistage cascade gas–liquid membrane contacting process.
membrane wetting, the appropriate number of modules should be
considered. The pressure drop ( P) of liquid flow in the lumen of the
fiber is determined using the Hagen–Poiseuille equation;
32vL ˛L
P =
2d
i
(24)
ˇ is contact angle and rP is average pore radius. The membrane
16
160
vl = 0.2 m/s
14
140
vl = 0.6 m/s
Maximum liquid-gas pressure
difference, ∆PL-G (kPa)
Total feed gas (l/min)
where ˛ is liquid viscosity, L is fiber length and di is inner diameter of the fiber. For the gas phase, the pressure drop is much lower
compared to that of liquid phase [11], thus, the pressure of the gas
side is assumed constant over the fiber length. In gas–liquid membrane contacting process, the pressure difference between liquid
and gas phases significantly affects the operation mode. The calculated maximum liquid–gas pressure difference (
PL–G = PL − PG )
is depicted in Fig. 10. It is found that pressure difference linearly
increases with number of module for all liquid velocities. Theoretically, the liquid can penetrate the membrane pores when the
pressure difference between liquid and gas sides is higher than
the wetting pressure. In contrast, the gas bubble formation would
occur if the pressure difference between gas and liquid phases
( PG–L = PG − PL ) is greater than bubbling pressure [23]. The wetting pressure or/and bubbling pressure ( PC ) can be determined by
Laplace equation;
−2
cos ˇ
PC =
(25)
r
P
where
is surface tension of the liquid absorbent (see Ref. [24]),
12
10
8
6
1.0 M
4
0.5 M
2
0.25 M
0
vl = 1.0 m/s
120
(∆Pc = 104.93 kPa)
100
80
60
40
20
0
0
1
2
3
4
5
6
7
Number of module
Fig. 9. The simulation results of the effect of number of module on total feed gas to
achieve 90%CO2 removal at different MEA concentrations (vl = 0.5 m s−1, yCO2 = 0.4).
0
1
2
3
4
5
6
Number of module
Fig. 10. The calculated maximum liquid–gas pressure difference,
number of modules and liquid velocities.
PL–G at different
183
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
(ii)
25
20
18
Non-wetted mode
Partially wetted mode
16
Partially wetted mode
Wetted mode
15
20
Non-wetted mode
10
5
% MEA consumption
% MEA consumption
(i)
14
Wetted mode
12
10
8
6
4
2
0
0
0.1
0.3
0.5
0.7
0.9
1.1
0.1
0.3
Concentration of MEA in feed liquid (mol/l)
(iii)
(iv)
20
0.9
1.1
35
30
16
14
Non-wetted mode
12
Partially wetted mode
% MEA consumption
% MEA consumption
0.7
Non-wetted mode
18
10
Wetted mode
8
0.5
Liquid velocity, v L (m/s)
6
Partially wetted mode
25
Wetted mode
20
15
10
4
5
2
0
0
0.15
0.25
0.35
0.45
0.55
0.65
Gas velocity, v G (m/s)
0
20
40
60
80
100
Concentration of CO 2 in feed gas (%Vol)
Fig. 11. The simulation results of total %MEA consumption of three–module cascade for three different operation modes including non-wetted, partially wetted (wetting ratio
estimated using Eq. (23)) and completely wetted modes: (i) At different MEA concentrations (vl = 0.2 m s−1 , vg = 0.5 m s−1 and yCO2 = 0.4); (ii) at different liquid veloc- ities (vg
= 0.5 m s−1 , yCO2 = 0.4 and CMEA = 0.25 M); (iii) at different gas velocities (vl = 0.2 m s−1 , yCO2 = 0.4 and CMEA = 0.25 M); (iv) at different gas compositions (vl = 0.2 m s−1, vg = 0.5 m
s−1 and CMEA = 0.25 M).
wetting results in performance deterioration, while the bubble formation causes gas loss during the operation [13,23]. Therefore, the
pressures of gas and liquid phases must be carefully controlled.
Generally, the pressure of liquid side is adjusted to be slightly
higher than that of gas side. The value of PL–G must be kept positive or equal to zero over the fiber length. If PL–G is fixed to be
zero at liquid outlet, the pressure of liquid at the inlet must be
adjusted to be greater than or equal to the value of pressure drop
of liquid flow over the fiber length. For the hollow fiber
membranes used in this work (rP = 0.08
m, ˇ = 93.4◦), MEA concentration of 0.25 M, the wetting pressure calculated from Eq. (25)
is 104.93 kPa. Therefore, the membrane wetting should not be
observed from our experiments although the operation was carried out at the conditions of highest wetting ratio (0.4 m s−1 and
CMEA 1.0 M). However, due to the distribution of membrane pore
size, the large pores may be wetted even the pressure difference (
PL–G) is lower than the wetting pressure estimated based on the
average pore size. Hence, the pressure drop over the length must
not exceed the wetting pressure. From Fig. 10 it can be seen that at
highest liquid velocity (1.0 m s−1), the systems with the num- ber
of module up to four modules show the maximum liquid–gas
pressure difference lower than critical wetting pressure. For fourmodule cascade, the pressure difference is about 92 kPa which is
quite close to the critical wetting pressure. To ensure that the
pressure drop would not significantly affect the membrane wetting, the pressure difference should be moderately lower than the
critical wetting pressure. Therefore, the three-module cascade
which shows the pressure drop around 66.9 kPa is selected for our
design.
MEA consumptions
Fig. 11 presents the simulation results of the total %MEA
consumption for three-module cascade at different operating
conditions of the three cases: non-wetted mode (x* = 0), partial wetted mode (x* estimated using Eq. (23)) and wetted mode
(x* = 1.0). The consumption of MEA is defined as the ratio of
moles of MEA consumed to the moles of MEA in feed ((CMEA,feed
− CMEA,out)/CMEA,feed × 100%). It can be obviously seen that the
system operated in non-wetted mode shows highest value of %MEA
consumption followed by partially wetted mode and completely
wetted mode, respectively. These reveal that the non- wetted mode
offers highest absorption performance resulted in highest MEA
consumption. For completely wetted mode, variation of MEA
concentration, gas velocity, liquid velocity as well as gas
composition slightly affects MEA consumption since the overall
resistance is dominated by the membrane phase.
Fig. 11(i) shows that at lower MEA concentrations, the percentages of MEA consumption are higher compared to those of higher
concentrations since the amount of MEA consumed in the system
with higher MEA concentration is higher than that of lower MEA
concentration due to higher rate of reaction. However, because of
the very high amount of MEA in the system with high MEA concentration, the MEA consumption of high MEA concentration is still
lower than that of lower MEA concentration. The effect of liquid
velocity on MEA consumption is depicted in Fig. 11(ii). At lower
liquid velocities, the percentages of MEA used for the reaction with
CO2 are higher compared to those of higher liquid velocities due to
longer contact time. Fig. 11(iii) reveals the %MEA consumption for
the system with different gas velocity. It is found that the %MEA
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S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Module 1
Module 3
Module 2
Fig. 12. Schematic diagram of three-module cascade membrane contactors with
different gas flow patterns (i) individual gas flow (G-ID), (ii) gas flow in series (G-IS).
CCO2,i/CCO2,feed in gas phase
1
0.8
0.6
0.4
0.2
G-ID
G-IS
consumption slightly changes with increasing gas velocity. As mentioned previously in Section 5.1, the gas phase velocity did not affect
the absorption flux, thus, the increase in gas velocity does not affect
%MEA consumption. Fig. 11(iv) presents the effect of CO2 concentration in gas phase on %MEA consumption. It can be seen that the
consumption of MEA increases with CO2 concentration due to the
enhanced driving force.
Total feed gas (l/min)
(i)
1
0
2
Fig. 13. Simulation results of the concentration profile of CO2 in gas phase of two
different gas flow patterns for achieving 90%CO2 removal.
However, the concentration of CO2 in gas phase considerably influences the absorption flux since the gas composition directly affects
the driving force of the system. Fig. 13 presents the concentration
profile of CO2 in gas phase along the length of three module cascade
for achieving CO2 removal of 90%. The values of CCO2,i/CCO2,feed at
the inlet and outlet are 1.0 and 0.1, respectively. For G-IS, the gas is
fed to the first module at z/L = 3 and leaves from the last mod- ule
at z/L = 0 (the gas flow direction is from the right to the left hand
side). It is known that the 3 module arrangement is equiv- alent to
one module with the length 3 times of a single module.
20
18
(ii)
16
9
8
14
12
10
8
6
4
G-ID
G-IS
2
7
6
5
4
3
2
G-I
1
D
G-IS
0
0
0
0.2
0.4
0.6
0.8
1
0
1.2
0.2
0.4
0.6
0.8
Concentration of MEA in feed liquid (mol/l)
35
(iv)
Total feed gas (l/min)
25
G-ID
30
1
Liquid velocity, v l
(m/s)
G-IS
Total feed gas (l/min)
(iii)
3
z/L1
Total feed gas (l/min)
Gas flow patterns
Fig. 12 shows two different gas flow patterns; (i) individual gas
flow (G-ID), the gas is fed separately into each module to achieve
the desired %CO2 removal at the outlet of each module, (ii) gas flow
in series (G-IS), which is similar to that in Fig. 8. The result from
Section 5.3 shows that the resistance in gas phase is only 3 and
6.5% for MEA concentration of 0.25 and 1.0 M, respectively. Therefore, the effect of gas velocity on absorption flux can be neglected.
0
25
20
15
G-I
D
G-IS
20
15
10
5
10
5
0
10
30
50
70
%Removal of CO2
90
110
10
30
50
70
90
110
Concentration of CO 2 in feed gas (%Vol.)
Fig. 14. Comparison of total feed gas of two different gas flow patterns: (i) at different MEA concentrations (vl = 0.5 m s−1 , yCO2 = 0.4 and CO2 removal = 90%) (ii) at various liquid
velocities (CMEA = 0.25 M, yCO2 = 0.4 and CO2 removal = 90%); (iii) at different %CO2 removal (vl = 0.5 m s−1 , CMEA = 0.25 M and yCO2 = 0.4); (iv) at different gas compositions (vl = 0.5
m s−1 , CMEA = 0.25 M and CO2 removal = 90%).
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Required MEA solution (l/min)
6
G-ID
5
G-IS
4
3
2
1
0
0
1
2
3
4
5
6
7
Total feed gas (l/min)
Fig. 15. Comparison of required MEA solution of two different gas flow patterns for
achieving 90% CO2 removal at various total feed gases.
Therefore, the CO2 concentration decreases continuously along the
length from z/L = 3 to z/L = 0 which is the same as that observed for
one module with module length of 3L (150 cm). When the gas mixture at high CO2 concentration is fed into Module 3, CO2 is partially
absorbed and the CO2 concentration is greatly reduced in this module. The reductions in CO2 concentration for Modules 2 and 1 are
less significant because the driving force for G-IS can be ranked as
Module 3 > Module 2 > Module 1. In case of G-ID, the gas flow direction is the same as G-IS. The gas is individually fed to one module
to achieve 90% removal at the outlet of each module. The values
of CCO2,i/CCO2,feed at the inlet and outlet for each module are the
185
same, 1.0 and 0.1, respectively. Therefore, the similar concentration
profile is observed.
Fig. 14(i) compares the total feed gas of two different gas flow
patterns at various MEA concentrations. The simulation result
reveals that G-ID gives higher performance compared to G-IS for
all MEA concentrations. The total feed gas of G-ID is higher than
that of G-IS roughly 17 and 33% for MEA concentrations of 0.125
and 1.0 M, respectively. This trend can be explained that at low
MEA concentrations, the percentage of MEA consumption in the
Module 1 is higher than that of high concentration. The absorption
rate significantly decreases in the next module. This results in the
slight reduction of CO2 concentration in gas phase for Module 3,
but the decrease is more significant for Module 1. Therefore, the
driving force of two gas flow patterns at lower MEA concentration
is closer compared to at high concentration.
Fig. 14(ii) illustrates that G-ID has higher system performance
than G-IS for all liquid velocities. For G-IS, the total feed gas
increases with liquid velocity. In case of G-ID, the total feed gas
sharply increases at low liquid velocities (0.1–0.2 m s −1), then,
there is a slight increase at moderate velocities (0.3–0.5 m s −1),
and a gradual decrease at high velocities (0.6–1 m s −1). These can
be explained by referring to the resistance analysis (Section 5.3).
At low liquid velocities, the mass transfer resistance is dominated
by liquid phase, thus, the system performance is considerably
enhanced with increasing velocity. For the moderate liquid velocities, the wetted membrane resistance becomes comparable to the
liquid phase resistance leading to a slight improve in absorp- tion
flux, whereas at high velocities, the system is controlled by the
wetted membrane resistance. Increasing in liquid velocity could
increase membrane wetting rather than improving the liquid phase
mass transfer, resulting in gradual absorption performance
deterioration. The performance comparison of two different gas
flow patterns at various % CO2 removals and gas compositions is
Fig. 16. Schematic diagrams of three-module cascade membrane contactor with different liquid flow patterns (i) liquid flow in series (L-IS), (ii) liquid flow in series with
splitting (L-ISS), (iii) liquid flow in series with recycle (L-ISR).
186
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
(i)
(ii)
9
8.5
8
Total feed gas (l/min)
8
Total feed gas (l/min)
9
8.5
7.5
7
6.5
6
5.5
7.5
7
6.5
6
5.5
5
5
4.5
4.5
4
4
0
0.2
0.4
0.6
0.8
0
1
0.2
Liquid velocity, v l (m/s)
0.4
0.6
0.8
1
Liquid velocity, v l (m/s)
Fig. 17. Simulation results of (i) effect of liquid velocity on total feed gas at different splitting ratios, (ii) effect of liquid velocity on total feed gas at different recycle ratios (CO2
removal = 90%, CMEA = 0.25 M, vl = 0.5 m s−1 and yCO2 = 0.4).
(i)
(ii)
9
Total feed gas (l/min)
8.5
Total feed gas (l/min)
21
19
8
7.5
7
6.5
6
5.5
17
15
13
11
9
5
7
4.5
5
4
0
0.2
0.4
0.6
0.8
Liquid velocity, v l (m/s)
1
0
0.2
0.4
0.6
0.8
1
Liquid velocity, v l (m/s)
Fig. 18. Comparison of total feed gas for three different liquid flow cascades (i) MEA concentration = 0.25 M, (ii) MEA concentration = 1.0 M, (CO2 removal = 90% and yCO2 = 0.4).
represented in Fig. 14(iii) and (iv), respectively. G-ID shows higher
system performance for all values of % CO2 removals and gas
compositions. The total feed gas into the system decreases with
increasing % CO2 removal and CO2 concentration in feed gas for
both gas flow patterns.
Fig. 15 depicts the comparison of MEA solution required for
achieving 90% CO2 removal at various total gas flow rates. It is
clearly seen that required MEA solution for G-IS is much higher
compared to that of G-ID. At total gas feed of 5.92 l/min, the required
MEA solution for G-IS is approximately 5.1 l/min, while that of G-ID
is only 1.5 l/min. These reveal that MEA solution is more effectively
used in case of G-ID. Therefore, the gas flow pattern of G-ID is
selected to study the effect of liquid flow pattern in the multistage
cascade contactors.
Liquid flow patterns
The schematic diagrams of three-module cascade with different liquid flow patterns are shown in Fig. 16. The first liquid flow
pattern is (i) Liquid flow in series (L-IS), which is similar to that in
Fig. 12(i). The second liquid flow pattern is (ii) Liquid flow in series
with splitting (L-ISS), from which the total liquid feed is partially
fed into Module 1 and the split liquid is partly combined with the
stream leaving Module 1 before entering Module 2. It is noted that
before entering module 3, the remained liquid is combined with the
liquid leaving Module 2. The third liquid flow pattern is (iii) Liquid
flow in series with recycle (L-ISR), from which the liquid feed flow
rate to each module is the same. After leaving Module 3, the liquid
with flow rate of QL,R is recycled and is combined with the original
feed before entering Module 1.
Fig. 17(i) depicts the effect of splitting ratio (˛) on system
performance of L-ISS at different liquid velocities (with MEA concentration of 0.25 M). It is found that L-ISS shows higher system
performance compared to L-IS (˛ = 0) at high liquid velocities,
whereas the splitting ratio of 0.3 gives the highest system performance. The total feed gas of L-ISS with splitting ratio of 0.3 is higher
than that of L-IS (˛ = 0) around 16.7% at liquid velocity 0.9 m s−1. The
effect of recycle ratio (ϕ ) on system performance for L-ISR at various liquid velocities (MEA concentration 0.25 M) is represented in
Fig. 17(ii). It is found that L-ISR can moderately improve system
performance at low liquid velocities. The Highest system performance is observed at recycle ratio of 1.0. The total feed gas of L-ISR
with recycle ratio of 1.0 is greater than that of L-IS (ϕ = 0) roughly
7.7% at liquid velocity 0.1 m s−1.
Fig. 18(i) and (ii) compares the system performance of three
different liquid flow patterns. For MEA concentration of 0.25 M, LISR can moderately improve the system performance at low liquid
velocities, whereas L-ISS considerably enhances the performance at
high liquid velocities. For MEA concentration of 1.0 M, the sys- tem
performance can be significantly increased by L-ISR at low to
moderate liquid velocities (0.1–0.5 m s−1). The total feed gases are
enhanced approximately 44.6 and 7.3% for liquid velocities of 0.1
and 0.4 m s−1, respectively. The L-ISS does not improve the system
performance at any liquid velocity. The reason for these differ- ent
results for typical MEA concentrations is due to the different
187
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
membrane wetting behaviors. At low MEA concentration, the values of x* greatly increase with liquid velocity (Section 5.2). At high
liquid velocities, the wetted membrane resistance controls the
overall resistance, and L-ISS can reduce membrane wetting ratio by
reducing the liquid flow rate in each module resulting in improv- ing
the performance. On the other hand, at low liquid velocities, the
overall resistance is dominated by liquid phase resistance, and L-ISR
can enhance the mass transfer coefficient of liquid phase by
increasing the liquid velocity in each module leading to enhanc- ing
absorption flux. In case of high MEA concentration, the wetting ratio
slightly increases with liquid velocity, and the liquid phase
resistance is comparable to wetted membrane. The reduction of
liquid flow in the module by splitting (L-ISS) does not considerably reduce the membrane wetting; thus, the system performance is
not enhanced for all liquid velocities. L-ISR can improve system
performance at low to moderate liquid velocities, because in this
range the liquid phase resistance still influences the system. The
increase of liquid velocity can improve the mass transfer in liquid
phase rather than increasing the membrane wetting. The results of
multistage design based on wetting ratio obtained from the experimental results reveal that different wetting behaviors at various
operating conditions (liquid velocity and MEA concentration) significantly affects the design of membrane contactors.
application. The operation of multistage cascade membrane contactor offers several advantages including the improved system
performance and the use of MEA effectively. However, the design
may be different depending on the conditions of each system.
Therefore, the experimental study is still necessary in order to
obtain the specific results for being used in the process design.
Acknowledgments
The authors gratefully acknowledge the financial support from
the Royal Golden Jubilee program and Senior Research Scholar
Grant from Thailand Research Fund (TRF) and from King Mongkut’s
University of Technology Thonburi (KMUTT).
Appendix A.
The diffusivity of CO2 in gas mixture (CO2–CH4), DCO2,G can be
calculated by the following [25]:
DCO2,G = 0.001858T
3/2(1/M
CO
2
P 2
AB
6. Conclusions
The experiments of CO2 absorption by MEA using PVDF hollow fiber membrane contactor were carried out. The mathematical
model
has
been
developed
to
predict
the
system
performance.
The experimental results were validated with the proposed model
for estimating the wetting ratio (x*) as function of liquid velocity
and MEA concentration. At low MEA concentration, the value of x*
greatly increases with liquid velocity, while that of high concentration slightly increases with velocity. The resistance analysis
demonstrates that for MEA concentration of 0.25 M, the
liquid
phase resistance dominates the overall resistance at low liquid
velocities. The wetted membrane resistance becomes comparable
to that of liquid phase at moderate velocities, and controls the
system at high liquid velocities. For MEA concentration of 1.0 M,
the wetted membrane resistance is comparable to the liquid phase
resistance for low to moderate liquid velocities, and dominates the
overall resistance at high velocities.
The suitable hollow fiber membrane module with effective fiber
length of 50 cm is selected for the design of multistage membrane
contactors. The three-module cascade is selected based on the pressure drop. The system performance is evaluated as the total feed
gas and MEA consumption. By comparing three different modes of
operation, the non-wetted mode presents highest MEA consumption followed by partially wetted mode and completely wetted
mode, respectively. Different gas and liquid flow patterns are com-
1/2
)
(A.1)
˝D
where MCO2 and MCH4 are the molecular weights of CO2 and CH4,
respectively. P is pressure in atm,
AB is characteristic length (m),
˝D is Collision integral. These parameters can be estimated by the
followings;
+
A
=
AB
B
(A.2)
2
where
A and
B are characteristic lengths of CO2 and CH4, respectively, and it can be determined as;
1.585Vb,A
=
A
(A.3)
1 + 1.3
2
where Vb is liquid molar volume at normal boiling point (cm3/mol).
The value of
can be estimated by the following;
=
1.94 × 10−3ı2
p
(A.4)
V bT b
where
p
is dipole moment (Debye). Tb is boiling temperature (K).
The Collision integral (˝D) is in function of temperature determined as;
a
˝D =
c
+
b
)
(∗ )∗
T
pared. For gas flow pattern, G-ID shows higher system performance
compared to G-IS for all operating conditions studied. In case of liquid flow pattern, L-ISS achieves the highest performance at splitting
ratio (˛) of 0.3, while that of L-ISR is obtained at recycle ratio (ϕ ) of
1.0. For low MEA concentration (0.25 M), the L-ISR can moderately
improve the system performance at low liquid velocities, while
L-ISS considerably enhances the performance at high liquid velocities. For the system with high MEA concentration (1.0 M), L-ISR
can improve the performance at low to moderate liquid velocities,
whereas the L-ISS does not improve the system performance at any
liquid velocity.
In conclusion, this work presents the design of multistage cascade membrane contacting process for chemical absorption of
CO2 that has never been reported in the literature. The process
design gives the insight and guideline for the scale up for real
+ 1/MCH4
exp(dT
+
e
exp(fT
∗)
g
+
exp(hT
(A.5)
∗)
where a, b, c, d, e, f, g and h are constants. T* can be calculated as the
following equation;
T
∗
=
ˇT
(A.6)
AB
where ˇ is Boltzmann’s constant and
AB
determined by;
AB
=(
A
B
is characteristic
energy
)1/2
(A.7)
where
respec-
A
and
B
are characteristic energy of CO2 and CH4,
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
tively.
188
S. Boributh et al. / Journal of Membrane Science 401–402 (2012) 175–189
Nomenclature
A
CA
CAG,in
CAGout,
CA,eq
CR,L
dh
Greek letters
(m2)
contact area
concentration of CO2 in bulk liquid (mol m−3)
inlet concentration of CO2 in gas phase (mol m−3)
outlet concentration of CO2 in gas phase (mol m−3)
concentration of CO2 at gas–liquid interface (mol
m−3)
concentration of absorbent (mol m−3)
hydraulic diameter (m), dh = d2 − n · d2/ds + n · d0
s
0
di
dint
dln
inner diameter of membrane (m)
interfacial diameter (m2)
logarithmic mean diameter of non-wetted membrane (m), dln = d0 − dint/ln(d0/dint)
d’ln
logarithmic mean diameter of wetted
membrane
(m), d∗l = dint − di /ln(dint /di )
n
d0
outer diameter
of membrane (m)
dpore
average pore size diameter of membrane (m)
ds
module diameter (m)
DG,eff
gas effective diffusivity (m2 s−1)
DG,Kn
Knudsen diffusion coefficient (m2 s−1)
DG,M
Fickian’s molecular diffusion coefficient (m2 s−1)
Di,G
diffusivity of component i in gas mixture (m2 s−1)
Di,L
diffusivity of component i in liquid solution (m2 s−1)
E
enhancement factor (dimensionless)
E∗
asymptotic infinite enhancement factor
∞
fA
FA
FI0
Gz
Ha*
H
JA,i
KL
kG
kL
kM
k’M
k2
L
Mi
n
N
P
PA
QG
189
molar flow rate of CO2 in liquid phase (kgmol s−1)
molar flow rate of CO2 in gas phase (kgmol s−1)
molar flow rate of inert component in gas
phase (kgmol s−1)
Graetz number (dimensionless)
Hatta number (dimensionless)
Henry’s constant (kPa m3 kgmol−1 )
absorption flux of CO2 (mol m−2 s−1)
overall mass transfer coefficient (m s−1)
gas mass transfer coefficient (kgmol m−2 kPa−1 s−1)
liquid mass transfer coefficient (m s−1)
gas-filled pores membrane mass transfer coefficient
(m s−1)
liquid-filled pores membrane mass transfer coefficient (m s−1)
second-order reaction rate constant (L mol−1 s−1)
length of hollow fiber membrane (m)
molecular weight of component i (kg kgmol−1 )
number of fibers
number of stages
total system pressure (kPa)
partial pressure of CO2 (kPa)
total volumetric flow rate of gas phase (m3 s−1)
QL
total volumetric flow rate of liquid phase (m3 s−1)
Re
Reynolds number (dimensionless)
RG
gas phase resistance (s m−1)
RL
liquid phase resistance (s m−1)
RM
gas-filled pores membrane resistance (s m−1)
R’M
liquid-filled pores membrane resistance (s m−1)
Rtot
total resistance (s m−1)
Sc
Schmidt number
(dimensionless) Sh
Sherwood number
(dimensionless) T
temperature (K)
x*
wetting ratio (dimensionless)
yCO2
molar fraction of CO2 in gas mixture (dimensionless)
˝D
ˇ
AB
ıdry
ıwetted
˛
ϕ
εM
R
M
i
collision integral for molecular diffusion (dimensionless)
Boltzmann’s constant (J K−1)
characteristic length (m)
density (kg m−3)
dry membrane thickness (m)
wetted membrane thickness (m)
split ratio (dimensionless)
recycle ratio (dimensionless)
membrane
porosity
(dimensionless)
kinematic viscosity (cm2 s−1)
stoichiometric coefficient of reaction
membrane tortuosity (dimensionless)1
characteristic energy of species i (J K− )
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