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www.sciencemag.org/content/361/6406/1008/suppl/DC1
Supplementary Material for
Zeolitic imidazolate framework membranes made by ligand-induced permselectivation
Xiaoli Ma*, Prashant Kumar, Nitish Mittal, Alexandra Khlyustova, Prodromos Daoutidis,
K. Andre Mkhoyan, Michael Tsapatsis*
*Corresponding author. Email: tsapatsis@umn.edu (M.T.); maxx0183@umn.edu (X.M.)
Published 7 September 2018, Science 361, 1008 (2017) DOI: 10.1126/science.aat4123
This PDF file includes:
Materials and Methods Supplementary Text Figs. S1 to S10 Tables S1 and S2 References
2
1. Materials and Methods
1.1 Preparation of porous supports
Porous α-alumina disk supports (22 mm in diameter) were fabricated by a colloidal processing
method reported before (41). α-alumina powder (CR-6, Baikowski, average particle size of 500
nm) was added to water to make a 50 wt% suspension. The pH of the suspension was adjusted to
~2.2 by 1M HNO3 solution. After horn sonication and degassing, approximately 4 g alumina
suspension was transferred into an annular polytetrafluoroethylene (PTFE) cylinder standing
vertically on a 0.2 µm nylon membrane (Whatman). The backside of the nylon membrane was
evacuated for 2 h under a vacuum of ~13 kPa to remove the water from the alumina suspension.
The as-formed alumina disk green body was dried overnight and then sintered at 1050 °C for 3 h
with a ramping rate of 2 °C/min. The mesoporous γ-alumina layer was prepared on the surface of
α-alumina support by a sol-gel method reported earlier (20). 1 M boehmite sol was prepared and
mixed with 3 wt% polyvinyl alcohol (PVA) solution to make a coating sol. The γ-alumina sol was
coated on the α-alumina surface by a slip-casting method. After drying in lab air overnight, the
support was sintered at 450 °C for 3 h with a heating and cooling rate of 0.5 °C/min.
1.2 All-vapor-phase membrane synthesis
The nanocomposite membranes were fabricated by an all-vapor-phase approach. ZnO was
deposited inside the pores of γ-alumina supports by an atomic layer deposition (ALD) process
(Savannah series from Cambridge NanoTech). The support was placed horizontally inside the
deposition chamber with the γ-alumina side facing upwards. The ALD was conducted at 125 °C
using diethylzinc and water as vapor precursors. A typical ALD cycle consisted of 0.015 s
exposure to water, followed by 5 s vacuum purge, 0.015 s exposure to diethylzinc, and 5 s vacuum
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purge. The number of deposition cycles was varied from 1 to 50 to change the amount of ZnO
deposited. The γ-alumina support after ZnO deposition was placed vertically inside a liner with
~0.2 g solid 2-methylimidazole (mIm) placed on the bottom. The distance between the support and
the solid mIm is ~ 4 cm. The liner was then sealed in an autoclave and heated at 125 °C for 24 h
for the conversion of the ALD-deposited ZnO to a porous zinc-imidazolate coordination
compound. After the growth, the membranes were activated by heat treatment under vacuum at
100 °C for 24 h or by methanol washing at room temperature for 4 h before gas permeation
measurements.
1.3 Water washing experiments
Certain membranes were washed with DI water to study the relationship between membrane
structure and separation performance (permeance and selectivity). The membrane was placed
inside a stainless-steel permeation cell. The front side (γ-alumina side) of the membrane was
washed with ~100 g flowing DI water with a flow rate of ~10 g/min. After washing, the membrane
was dried in an oven at 70 °C under N2 flow (100 mL/min) for 18 h before gas permeation
measurements and structural characterization.
1.4 Gas permeation measurements
Gas permeation/separation measurements were performed in a home-built constant volume
variable pressure apparatus. The membrane was tightly sealed with Viton o-rings inside a stainless-
steel permeation cell. The cell was loaded in the permeation apparatus, followed by a 2 h
evacuation of the whole system before each gas measurement. The feed pressure can be controlled
by a pressure regulator. The rate of pressure change on the permeate side was used to calculate the
4
gas permeance. For single gas permeation tests, the ideal selectivity/separation factor was
determined by the ratio of permeances of the two pure compounds. For mixed gas measurement,
an equimolar C3H6/C3H8 mixture (Fig. 1, Figs. S1 and S2) and C3H6/C3H8 mixtures with a range
of compositions (Fig. S3) at a total flow rate of 100-200 mL/min was provided on the feed side,
and the permeate side was either swept with 120-200 mL/min N2 (Fig. 1 and Fig. S1) or was under
vacuum (Figs. S2A and S3) or was kept at 1 atm undiluted (i.e., no sweep gas) permeate (Fig.
S2B). The gas composition was measured by gas chromatography (GC) equipped with a flame
ionization detector (FID) detector and a capillary column. The separation factor was defined as the
molar ratio of propylene/propane in the permeate divided by their molar ratio in the feed. To
evaluate the membrane performance under different pressures, the feed pressure was varied from
1 to ~7 atm by a regulator on the retentate line.
1.5 Characterization
X-ray diffraction (XRD) patterns were recorded using a Bruker-AXS (Siemens) D5005
diffractometer with a CuKα (λ = 0.15406 nm) radiation source. Scanning electron microscopy
(SEM) images were acquired using a Hitachi SU8230 scanning electron microscope at an
accelerating voltage of 5 kV. SEM specimens of supports and membranes were coated with 5 nm
of Ir.
1.6 TEM
Cross-sectional membrane specimens (<50 nm in thickness) for TEM imaging were prepared by a
focused ion beam (FEI DualBeam Helios G4) using 30 kV Ga ions, followed by 5 kV and 2 kV
ion milling to remove damaged surface layers created due to heavy Ga ion impingement.
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Furthermore, polishing by milder ion milling at 1 kV at tilts of +/- 5o were performed to reduce
any structural damage to the cross-sectional specimen.
High-angle annular dark field – scanning transmission electron microscopy (HAADF-STEM) data
were obtained using an aberration-corrected FEI Titan G2 60-300 STEM at 60 kV, 10-30 pA beam
current, 24 mrad convergence angle with ~50 mrad ADF-detector inner angle. The electron probe
was corrected using a CEOS-DCOR probe corrector. Energy-dispersive X-ray (EDX) spectrum
imaging was performed using a Super-X system at 60 kV and 10-30 pA beam current with 24 mrad
convergence angle. Frame-by-frame spatial drift correction was enabled using Bruker Espirit 1.9
software.
2. Process-scale Assessment
In this section, a process-scale assessment of membrane separation is performed and compared
to distillation, which is the current industrial practice for propylene-propane separation. The main
conclusions from the assessment are summarized below while detailed analysis is described in the
sub-sections.
A first look at a distillation column achieving an annual production of 250,000 tons of
propylene reveals an energy requirement of ~40 MW, reflecting an energy intensive process (42).
However, this is an overestimation of the required energy because distillation columns can be heat-
integrated (43, 44). Indeed, our calculations show that heat-pump based heat-integrated columns
(also known as vapor recompression design) (Fig. S6B) are capable of achieving up to 75% energy
savings. However, such designs are not routinely employed due to operational and control
challenges (45, 46). Instead, heat-integration can be implemented through other parts of the plant
generating extra amount of low-grade heat, such as, quench water. Therefore, a range of possible
6
scenaria for the level of heat integration should be examined with thermally non-integrated and
fully-integrated distillation columns as the upper and lower bounds of energy requirement.
While it is difficult to replace distillation by membranes (Fig. S7), a hybrid membrane-
distillation combination (Fig. S6C) is attractive for debottlenecking of existing distillation columns
(Fig. S8). Using the currently demonstrated membrane performance of ~100 GPU and selectivity
of ~50 for a 7 bar feed and 1 bar permeate (Fig. S2 and S3), a 35% increase in productivity can be
attained as shown in Fig. S8. Even for a 50% heat-integrated column (i.e., with 50% of the reboiler
duty provided by heat-integration), our membranes can achieve ~25% savings in energy
requirement over heat-integrated distillation (Fig. S9A). This corresponds to a savings of 0.17
cents/kg of propylene. For reference, the operational cost of a 50% heat-integrated column is 1.1
cents/kg and the cost of polymer-grade propylene is 65 cents/kg (47). Considering the current
annual production of propylene at 50 million tonnes (1, 32), a membrane-distillation hybrid can
save ~100 million USD annually, if fully implemented. If the selectivity of the membranes could
be improved to 200, then energy savings for a 50% heat-integrated column would increase to ~40%
(Fig. S9B).
The membrane area required to achieve production of 250,000 tonnes/year of propylene can
be estimated based on the expected membrane permeance at operation conditions. To be on the
conservative side, using the achieved permeance of 100 GPU and selectivity of 50 at feed pressure
of 7 bar, the area requirement is calculated to be ~12,300 m2 for a plant producing 250,000
tonnes/year of propylene (Fig. S9C). A breakeven in capital cost can be achieved at a membrane
cost of ~$130/m2.
A well-established heuristic guideline (48, 49) is that for optimal use of a selective membrane,
the ratio of feed (retentate) to permeate pressures (often called the pressure ratio) should not be
7
much smaller than the selectivity. Although we have experimentally achieved membranes with
selectivity of ~50, in Fig. S9C we evaluated them with a relative small pressure ratio of 7, just to
be on the conservative side. Since our data (Fig. S2) indicate that flux continues to increase
between 5 and 7 bar (consistently with a constant permeance of 100 GPU up to the highest tested
feed pressure), we can also consider a more positive scenario of constant permeance up to 15 bar
feed. An even more attractive case can then be made for the use of our membranes at this, higher,
pressure ratio (Fig. S8 and Fig. S9E). The required area to achieve the same capacity increment
and energy savings now drops to 4,800 m2 (Fig. S8) and the breakeven in capital cost can be
achieved at a membrane cost of ~$330/m2 (Fig. S9E). Even more favorable capital cost savings
can be achieved at improved membrane selectivity of 200 at the pressure ratio of 7 and 15 (Fig.
S9D and S9F).
The large membrane area requirement and low cost can only be achieved with a highly scalable
membrane production process, like the liquid-free LIPS process.
Another attractive potential use of the propylene selective membranes, is the recovery of
propylene that is lost in the purging process in polymerization plants, which is estimated in excess
of 5,000,000 lb with a corresponding value of $1 million, per year, per polymerization plant (9,
50, 51). Although purge is a small fraction of the plant capacity, considering the polypropylene
production of more than 50 million tonnes/year (52), the lost propylene is worth more than 200
million USD annually. Conservative estimations using propylene permeance of 100 GPU and a
selectivity of 5, indicate a reasonable membrane area requirement of ~250 m2 and pay-back periods
smaller than 1 and 3 years (Fig. S10), for membrane costs of $1,000 and $10,000/m2, respectively.
Compared to C3 splitter, the reactor purge application employs much smaller (100-fold less) area
8
and can correspondingly incur higher cost. Thus, membranes for reactor purge application can be
industrialized at current membrane performance.
The details of the distillation and membrane process separation are discussed next.
2.1 Separation using distillation
2.1.1. Conventional distillation column
C1-C4 hydrocarbons are usually obtained via cracking and then separated using a series of
distillation columns (53). The last step in the production of propylene is propylene/propane
separation and is one of the largest energy consuming separation steps. The feed composition for
the separation, based on cracking product composition, is ~70% propylene and ~30% propane. A
similar feed composition has been considered in other studies (42, 53–56) and, thus, is also selected
in this study for process-scale assessment. A distillation column (Fig. S6A) for obtaining 99.7 mol%
pure propylene (polymer-grade propylene) at an annual production of 250,000 tonnes and a 99
mol% propane as bottoms was simulated using RadFrac model with Redlick-Kwong-Soave
property-set in Aspen Plus. The number of trays was fixed at 247, and a uniform pressure of 15
bar (top temperature of 35 °C) is considered, which allows cooling water to be used as the cooling
duty. As shown in Table S2, ~40 MW of reboiler and condenser duties are required reflecting an
energy-intensive process.
2.1.2. Heat integrated distillation column
A heat integrated column was also modeled. The heat integration can be broadly performed in
two ways: (i) in external heat integration, the vapor from the top tray is compressed to an extent
such that the increase in temperature allows heat-exchange with the condenser before entering as
a reflux stream, and (ii) in an internally heat integrated column, additional heat exchange can also
9
take place between the trays in the stripping and rectification sections of the column (57). It has
been shown that an externally heat integrated column consumes ~6-fold less energy when
compared to a conventional column for propylene/propane separation (58, 59). Further, an
internally heat integrated column only achieves marginal or no energy savings over an externally
heat integrated column for propylene/propane separation, and adds more design complexities (59,
60). Thus, an externally heat-integrated column, also known as vapor recompression column, is
selected in this study (Fig. S6B). In a conventional distillation column, the condenser temperature
is lower than the reboiler temperature and prohibits heat integration. In a vapor recompression
column, the vapor stream is compressed so that its temperature rises above the bottom stream
temperature and enables heat integration between condenser and reboiler. It should be noted that
the smaller the difference in the boiling points of the components, the smaller is the energy required
for compression to enable heat-integration.
As the heat of vaporization decreases with temperature, compressing to higher pressures (thus,
higher temperature) results in lesser total heat recovery on condensation at higher temperature.
Thus, very high pressure (or temperature) is required so that large total sensible heat can also be
used to fully heat-integrate the reboiler and the condenser. This results in higher compression work.
We found that compression to 30 bar is required to fully heat-integrate the column, which requires
6.5 MW and 4.7 MW at minimum approach temperature of 5 and 10 °C, respectively. This is also
consistent with the 6-fold lower energy consumption reported in the literature (58, 59). We also
considered a case where the distillate was compressed to only 20 bar and the rest of the reboiler
energy requirements were fulfilled by external steam. This resulted in 2 MW of compressor work
and 4 MW of steam at minimum approach temperature of 5 °C. Although the total energy, ~6 MW,
is larger than that of fully heat-integrated case (4.7 MW), the fuel-equivalent energy (11 MW fuel-
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equivalent) is lower than that of fully heat-integrated case (15.5 MW fuel-equivalent). Thus, this
partial heat-integrated column, which results in ~75% energy savings is considered as base case
for heat integrated column.
2.2 Separation using membranes
A counter-current membrane model with a constant permeance and selectivity, and a plug flow
for both the retentate and the permeate side is considered. A uniform total pressure was assumed
on both sides. The governing flow equations for each component are as follows:
where the positive sign refers to the retentate side while the negative sign refers to the permeate
side, F is the flow rate along the membrane, J is the flux through the membrane, and r is the radius
of the tube. The flux through the membrane is given by:
where П refers to the permeance, p refers to the partial pressure, and sel refers to the selectivity,
while ‘ret’ and ‘perm’ denote the retentate and permeate side, respectively.
Transport rates through membranes are determined by adsorption-diffusion phenomena, which
are pressure-dependent. In this study, the experimentally determined performance was obtained at
a feed pressure of up to 7 bar and total permeate pressures of less than or equal to 1 bar. Since the
amount adsorbed will reach saturation at high enough pressures, the driving force for membrane
transport ceases to increase linearly (and eventually completely) with feed pressure beyond a
certain value. Further, the driving force becomes negligible if the permeate pressure is close to or
higher than the saturation pressure, and the membrane becomes non-selective. While in most of
the membrane modeling studies in the literature (54–56, 61), a permeate pressure of 3-5 bar is
ii π2 rJ
dx
dF
(1)
jiji
permiretii
Sel
ppJ
,
,,i ;
(2)
11
considered, resulting in lesser energy requirement for permeate recompression to feed pressure
(~15 bar), the performance of membranes is not reported often in the open literature at such
conditions, and it is likely (due to reaching small loading gradients across the membrane and/or
reduced diffusivity at high loadings) that membranes will exhibit reduced actual performance than
the one used in simulations.
In this study a conservative approach is used and the total permeate pressure in the model is
maintained at 1 bar or less to correspond to the highest permeate pressure tested experimentally.
Further, to demonstrate the effect of pressure ratio (ratio of feed to permeate pressures), we
examine several feed pressures ranging from 5 bar to 15 bar. We note that our data (Fig. S2 and
S3) for membrane performance validate the use 100 GPU permeance and selectivity of 50 for up
to 7 atm feed pressure. Propylene purity against recovery obtained using a single-stage membrane
model at several selectivity values (2 to 500) for a 70 mol% feed at a total pressure of 7 bar on the
retentate side and a total permeate pressure of 1 bar are shown in Fig. S7A. Even with a selectivity
of 500, the polymer grade propylene (99.7 mol% purity) can be obtained only up to a recovery of
80% (as compared to 99.6% for distillation). Thus, high selectivity (greater than 500) membranes
are required to achieve the purity and recovery targets in single-stage membrane process.
Next, we set the recovery to 90% and investigate the effect of feed and permeate pressure
(pressure ratio) on the obtained purity and the required surface area. We also assess the importance
of the achieved selectivity of ~50, by comparing obtained purity and required area at lower and
higher selectivities.
For the currently established (experimentally validated) membrane performance (100 GPU,
selectivity of ~50, 7 bar feed, and 1 bar permeate) the area required for achieving 96 mol% purity
and 90% recovery is determined to be ~24,000 m2, while if we extrapolate the same performance
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(100 GPU, selectivity of ~50) to 15 bar feed, a higher pressure ratio enables a significant reduction
of the required area to ~9,000 m2 (Fig. S7B). If we consider the 30-fold higher permeance used by
Gottschlich and Roberts (62), their much smaller surface area estimation is in qualitative
agreement with our estimation. However, as stated earlier, membranes cannot achieve the required
propylene purity of 99.7 mol%, and require higher selectivity, which is also in agreement with
Gottschlich and Roberts (62).
The effect of selectivity at different pressure ratios (achieved using 5, 7, 10 and 15 bar feed
with 1 bar permeate; and 7 bar feed with different levels of vacuum in the permeate) are shown in
Fig. S7C and S7D. There is an increase in the achieved purity as the selectivity increases. With the
set 90% recovery, for a selectivity increase from 2 to 50, there is an increase in purity from 75
mol% to ~96 mol%, while purity gains become smaller as the selectivity increases above 50.
Moreover, as the pressure ratio increases, there is a relative small gain in purity accompanied with
a significant reduction in the required surface area (e.g., for selectivity of 50, a four-fold reduction
of required membrane area from pressure ratio of 5 to 15). The effect of different levels of vacuum
on purity for a total feed pressure of 7 bar on the retentate side and at a fixed selectivity of 50 is
also shown in Fig. S7C and S7D. Fig. S7C shows that small gains in purity are obtained as the
permeate pressure decreases below 1 bar. Fig. S7D includes the area requirement for a feed
pressure of 7 bar and permeate pressure of 0.5 bar, showing smaller area requirement than 7 bar
feed and 1 bar permeate, as anticipated by the increasing driving force between these two cases.
For the case of 90% recovery and the range of parameter values examined in Fig. S7C and
S7D, it is found that increasing selectivity has more pronounced effect than operating with vacuum
on the permeate side or increasing pressure on the retentate side. With a pressure ratio of 7 and
above, 90% recovery can be achieved with increasing purity as the selectivity increases without
13
significantly increased requirements for surface area. This suggests that, under these operating
conditions, the membranes are operating in the selectivity-limited regime as opposed to the
pressure ratio-limited regime (48, 49). However, if the pressure ratio is 5, then the purity gains
achieved by high increases in selectivity are accompanied with significant increase in area
requirements (Fig. S7D, red line).
We can conclude that the performance of our membranes is not sufficient for replacing
distillation. Much higher selectivities and ability to retain them at high pressure ratios will be
required to meet the purity and recovery achieved by distillation. For example, to achieve purity
and recovery of ~99% in a single stage membrane, at a pressure ratio of 15 (15 bar feed and 1 bar
permeate), a selectivity of 1,000 and the corresponding large area of ~50,000 m2 is required. Next,
a hybrid membrane-distillation process is evaluated for obtaining polymer-grade 99.7 mol%
propylene.
2.3 Separation using hybrid membrane-distillation
Several configurations have been proposed for hybrid membrane-distillation and can be
classified in two broad categories – (i) pre-treatment by membranes, which includes series and
parallel configurations, and (ii) polishing step by membranes, which includes top and bottom
configurations (62–66). Separation performance comparison among these configurations for
olefin/paraffin separation suggests that the optimal configuration depends upon several variables
including operation conditions and objective function (62, 65, 67, 68). In this study, only a series
configuration is analyzed, as also considered for other debottlenecking/retrofitting studies (9, 32,
69, 70), but other configuration are also worth investigating. As shown in Fig. S6C, in a series
configuration, the feed is directly sent to the membrane while distillation is used to achieve the
14
final product specifications. The stage location of retentate and permeate streams to the distillation
column are optimized to achieve minimal energy requirements for the same number of total trays
(= 247). A membrane stage-cut of 0.5 and selectivity values of 50 and 200 are considered. The
membrane stage cut is defined as the ratio of propylene (preferentially permeating component)
flow rate in the permeate stream to that in the feed, and it is equal to the recovery. As membranes
perform a part of the separation, the separation load on distillation is decreased leading to reduction
in reflux and boilup ratio. To maintain the same vapor and liquid flowrate within the column, i.e.,
to use the same column diameter, the capacity of the existing column (as it would be the case in
debottlenecking of an existing plant) would be increased by implementing hybrid membrane-
distillation. The increase in capacity for different selectivity values are shown in Fig. S8 which
suggests ~1.35-fold increment for selectivity of 50.
The operational energy requirements and capital expenses are also analyzed. While the
reduction in reboiler duty reduces the energy requirement, the permeate compression results in
additional operating expenses. The operating energy and savings varying with increment in heat-
integration are shown for a selectivity of 50 (Fig. S9A), and 200 (Fig. S9B). These results suggest
that hybrid configuration results in net energy savings even for columns with heat integration. The
additional capital cost (comprised of the membrane and the compressor costs) for capacity
increment is also calculated for two pressure ratios (7 and 15, corresponding to 7 and 15 bar feed
and 1 bar permeate) and compared to the capital expenses for base-case distillation. The capital
cost of base case distillation achieving an annual production of 250,000 tons of propylene is 7
million USD (2.80 cents/kg) (43). The capital expenses and savings for increased capacity in the
hybrid configuration are shown for several combinations of permeance and membrane cost at a
selectivity of 50 (Fig. S9C for pressure ratio 7 and S9E for pressure ratio 15), and 200 (Fig. S9D
15
for pressure ratio 7 and S9F for pressure ratio 15). For pressure ratio of 7, a membrane cost of
~$130/m2 is required to break even with distillation-only capital cost at the achieved levels of
selectivity of 50 and permeance of 100 GPU. Assuming that similar permeance (100 GPU) and
selectivity (~50) can be maintained at feed pressure of 15 bar, capital savings of 25% can be
achieved at a membrane cost of $200/m2. These results suggest significant potential of using the
currently reported membranes even with heat-integrated distillation column. They also highlight
the importance of being able to achieve selective membrane performance under high enough
pressure ratios.
2.4 Recovery from reactor purge
Propylene is a major precursor in the synthesis of important products, such as polypropylene,
cumene, isopropyl alcohol, etc. In such applications, a stream of 70-90 mol% propylene (shown
as separator recycle in Fig. S10A) is normally recycled back to the reactor (9, 50, 51). To prevent
build-up of propane, a part of this stream is also purged (shown as separator purge) and results in
propylene losses. These purge streams are usually lost or flared which is inefficient from both the
energy and the environmental considerations. Although separator purge is a small fraction of the
feed (and/or separator recycle stream), it contains substantial amount of propylene (~5 million lb
per polymerization plant annually) and can lead to considerable profits if recovered even up to
90%. The membrane separation has been proposed as a potential solution to recover this propylene
and recycle it back to the reactor. Although the feed to the reactor is polymer-grade (~ 99.7 mol%),
80-95 mol% permeate recycle has been found to be adequate (50, 51). This is because the separator
recycle stream is 70-90 mol% and thus the overall composition in the reactor is lower than the
polymer grade feed. Further, the permeate recycle is only a small fraction of the total feed (fresh
16
feed + separator recycle + permeate recycle) and only has a marginal effect on the total
composition.
As moderate purity is adequate for this application, separation by membranes can be performed
in a single-stage at reasonable selectivity. The results obtained for propylene purity against
recovery for several values of selectivity, and propylene purity against selectivity for several
values of feed composition are shown in Fig. S10B and S10C, respectively. As shown, even for a
70 mol% feed, 90% propylene can be recovered at > 80 mol% purity for selectivity as low as 5.
The membrane area required for 80 mol% propylene feed at 2.6 mol/s (equivalent to 5 million lb
propylene annually at a recovery of 90%) is shown in Fig. S10D. As shown, for a permeance of
100 GPU, membrane area of ~250 m2 will be required for a typical propylene polymerization
reactor.
An economic analysis is also carried out to evaluate the profits associated with purge stream
recovery. As the reactor is operated at high pressure (5 – 30 bar) (71–73), compression of the
permeate stream will constitute the majority of the operational cost. The major capital cost will
comprise the cost of the compressor and the membrane. This cost, then, shall be compared to the
value of recovered olefin to evaluate the scope of membranes for this application. Considering a
price of 20 cents/lb for refinery-grade propylene (47, 50), the total annual revenue for a recovery
of 5 million lb propylene per year amounts to $ 1 million. Correspondingly, three compressors of
< 10 kW each are required for permeate compression to 30 bar. Considering a cost of 7 cents/kWh
of electrical energy, this amounts to $17,000 in operational costs which is < 2% of the total revenue.
The corresponding total capital cost for compressor and membrane amounts to $30,000 and
$125,000 (considering installed membrane cost of $500/m2), respectively. This results in a
payback of < 1 year with > 5-fold return in the first year itself. A net present value analysis (NPV)
17
is also performed and the NPV profits are calculated at a discount rate of 10%. Furthermore, the
membrane was assumed to have a lifetime of 3 years and to cost 50% of the initial investment
when replaced. NPV profits over a period of 10 years are shown in Fig. S10E, which suggests
significant benefits of using membranes for reactor purge application.
18
Fig. S1. On-stream stability test of C3H6/C3H8 mixture separation performance. The
membrane tested was prepared from 10 cycles ZnO ALD. An equimolar C3H6/C3H8 mixture at a
total flow rate of 100 mL/min was used on the feed side, and the permeate side was swept with
120 mL/min N2. The feed pressure was varied from 1 atm to ~7 atm. The test was conducted at
room temperature.
19
Fig. S2. Effect of feed pressure on C3H6/C3H8 mixture separation performance. An equimolar
C3H6/C3H8 mixture at a total flow rate of 200 mL/min was used on the feed side. The permeate
side was (A) under vacuum; and (B) kept at 1 atm undiluted (i.e., no sweep gas) permeate. The
test was conducted at room temperature.
20
Fig. S3. Effect of feed gas composition on C3H6/C3H8 mixture separation performance. A
C3H6/C3H8 mixture with a total feed pressure of ~ 7 atm and a total flow rate of 200 mL/min was
used on the feed side. Vacuum was used on the permeate side. The test was conducted at room
temperature.
21
Fig. S4. High-magnification top-view SEM image of (A) γ-alumina support (B) γ-alumina support
with 10 cycles ZnO ALD before and (C) after ligand vapor treatment. The scale bars in A, B and
C are 100 nm.
22
Fig. S5. Spatial distribution of nitrogen across membrane cross-section. (A) ADF-STEM
image of the top surface of the Al2O3 support, ALD-ZnO and ZIF-membrane. (B) Averaged line
scan of STEM-EDX maps for nitrogen signal (N-K edge) obtained from the corresponding ADF-
STEM images shown in A along the marked arrows. (C) ADF-STEM image of the γ-alumina/α-
alumina interface for Al2O3 support, ALD-ZnO and ZIF-membrane. (D) Averaged line scan of
23
STEM-EDX maps for nitrogen signal (N-K edge) obtained from the corresponding ADF-STEM
images shown in C along the marked arrows. Scale bars in A from left to right are 60 nm, 60 nm
and 40 nm. Scale bars in C are 700 nm.
24
Fig. S6. Distillation and membrane-distillation configurations considered. Schematic of (A)
a conventional distillation column, (B) a heat-integrated vapor recompression distillation
column, and (C) a hybrid membrane-distillation column for propylene-propane separation.
25
Fig. S7. Calculations for single-stage membrane replacing distillation. (A) propylene purity
obtained for increasing recovery obtained using a single-stage membrane model for 70.0 mol%
propylene feed at a total pressure of 7 bar on the retentate side and a total pressure of 1 bar on
the permeate side, (B) membrane area required for increasing recovery at a permeance of 100
GPU and selectivity of 50 obtained using a single-stage membrane model for 70.0 mol%
propylene feed at several values of total pressure on the retentate side and total pressure of 1 bar
on the permeate side, (C) propylene purity obtained for increasing selectivity at a recovery of
0.9 for 70.0 mol% propylene feed at several values of total pressure on the retentate side and
total pressure of 1 bar on the permeate side; also shown (in black) is propylene purity obtained
for different levels of vacuum permeate at a recovery of 0.9 and a selectivity of 50 for 70.0 mol%
26
propylene feed at a total pressure of 7 bar on the retentate side, and (D) membrane area required
against selectivity for a recovery of 90% at a permeance of 100 GPU obtained using a single-
stage membrane model for 70.0 mol% propylene feed at several values of total pressure on the
retentate side and total pressure of 1 bar on the permeate side; also shown (in black) the case for
a total pressure of 7 bar on the retentate side and 0.5 bar vacuum on the permeate side.
27
Fig. S8. Membrane area required and capacity increment for propylene-propane separation
obtained by membrane-distillation hybrid over distillation column shown for selectivity of 50
and 200, and a total pressure of 7 and 15 bar on the retentate side and a total pressure of 1 bar
on the permeate side.
28
29
Fig. S9. Operational energy requirement and capital cost expenses and savings for
propylene-propane separation by membrane-distillation hybrid over that for heat-
integrated column. The heat-integrated fraction represents the fraction of distillation column
energy requirement assumed to be provided by heat-integration; the higher end of 0.75
corresponds to the fraction that can be achieved by full heat integration by a heat-pump. Feed
conditions and separation targets are provided in Table S2. Energy requirements and savings are
shown for a selectivity of (A) 50 and (B) 200. Capital expenses and savings are shown for a total
pressure of (C, D) 7 bar and (E, F) 15 bar on the retentate side and a total pressure of 1 bar on
the permeate side, with permeance of 100 and 200 GPU, and a selectivity of (C and E) 50, and
(D and F) 200. The capital cost for distillation in C-E are based on a column with no heat-
integration.
30
31
Fig. S10. Calculations for membrane retrofitting in a polymerization reactor purge stream.
(A) Schematic showing the original reactor configuration with solid lines and the membrane
retrofitting part with dashed lines. (B) Propylene purity against recovery obtained using a single-
stage membrane model for 80.0 mol% propylene feed at several values of selectivity. (C)
Propylene purity against selectivity obtained using a single-stage membrane model for 90.0%
recovery shown for several feed compositions. (D) Membrane area required for 80 mol%
propylene feed at a total flowrate of 2.6 mol/s. For a recovery of 90%, this would correspond to
5 million lb propylene annually. (E) Net present value of profits for 90% propylene recovery
from a reactor purge stream considering reactor pressure of 30 bar, annual propylene recovery
of 5 million lb, and membrane permeance and selectivity of 100 GPU and 5, respectively.
Membrane performance in B, C and D is simulated with a total pressure of 5 bar on the retentate
side and a total pressure of 1 bar on the permeate side.
32
Table S1.
Separation performance data, membrane preparation and gas permeation/separation
measurement conditions of the membranes shown in Figure 1D and 1E. All the mixed gas
permeation measurements used a 50:50 propylene/propane mixture in the feed side. The fluxes
were calculated from permeances and transmembrane pressure drop reported.
Substrate
Membrane
fabrication
method
Gas
permeation
method
Total
feed
pressure
(kPa)
Total
permeate
pressure
(kPa)
Permeate side
condition
C3H6
Permeance
(×10-10
mol/m-2s-1Pa)
C3H6 flux
(×10-4
mol/m-2s-1)
Separation
factor
Symbols
in
Figure
Reference
α-alumina Secondary
growth
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min He
285 14.4 34
(5)
277 14.0 35
245 12.4 31
206 10.4 45
378 19.2 28
α-alumina Secondary
growth
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar
78 3.95 89
(23)
83 4.2 63
156 7.9 50
110 5.57 75
167 8.46 44
α-alumina Secondary
growth
Wicke–
Kallenbach
~101
~0 Sweeping, 100
mL/min Ar
64 3.26 51
(33)
~101 55 2.80 50
~101 77 3.92 47
~101 62 3.14 61
~200 54 5.47 27
~400 43 8.74 14
α-alumina Secondary
growth
Wicke–
Kallenbach
~101
~0 Sweeping, 100
mL/min N2
110 5.57 30
(24) ~200 89.4 9.07 28.9
~300 76.4 11.8 28.1
~460 60.6 14.0 26.9
α-alumina Counter-
diffusion Time-lag ~100 ~0 Vacuum 25.0 1.25 59
(25)
α-alumina Counter-
diffusion Time-lag
~100
~0 Vacuum
11.0 0.55 135 (34)
33.0 1.65 113
α-alumina Counter-
diffusion Time-lag
~100
~0 Vacuum
120 6.00 20
(35) 52.0 2.60 7.2
390 19.5 6.9
α-alumina Counter-
diffusion Time-lag ~100 ~0 Vacuum 22.0 1.10 10
(26)
α-alumina Counter-
diffusion Time-lag ~100 ~0 Vacuum 70.0 3.50 42 (36)
α-alumina In-situ
Constant
pressure
variable
volume
~200 ~100 1 atm of
permeated gas
86.8 4.34 36
(27) 149 7.45 13.1
α-alumina Heteroepitaxial
growth
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar
461 23.3 84.8
(6) 370 18.7 209.1
309 15.7 163.2
33
Table S1 (continued).
Separation performance data, membrane preparation and gas permeation/separation
measurement conditions of the membranes shown in Figure 1D and 1E. All the mixed gas
permeation measurements used a 50:50 propylene/propane mixture in the feed side. The fluxes
were calculated from permeances and transmembrane pressure drop reported.
Substrate
Membrane
fabrication
method
Gas
permeation
method
Total
feed
pressure
(kPa)
Total
permeate
pressure
(kPa)
Permeate side
condition
C3H6
Permeance
(×10-10
mol/m-2s-1Pa)
C3H6 flux
(×10-4
mol/m-2s-1)
Separation
factor
Symbols
in
Figure
Reference
α-alumina In situ counter
diffusion
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar 213 10.8 50
(37)
α-alumina
Seeding+vapor
-phase
Ripening
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar 125 6.33 120 (18)
α-alumina
Microwave-
assisted
seeding and
secondary
growth
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar
208 10.5 40.43
(38)
144 7.28 30.77
α-alumina In situ counter
diffusion
Wicke–
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar
269 13.6 70.6
(39)
538 27.2 22.4
395 20.0 28
268 13.6 38
308 15.6 33.1
195 9.88 39.7
α-alumina
Postsynthetic
linker
exchange
Wicke-
Kallenbach ~101 ~0
Sweeping, 100
mL/min Ar 780 39.5 40 (10)
Polymer
Interfacial
Microfluidic
Processing
Wicke–
Kallenbach
~100
~0
Sweeping, Ar
151 7.92 184.4
(40) ~400 112 22.8 176.2
~700 98.7 34.7 135
~850 91.5 39.1 90.3
Polymer
Interfacial
Microfluidic
Processing
Wicke–
Kallenbach ~101 ~0
Sweeping, 30
mL/min Ar
221 11.2 65
(28) 114 5.8 24
Polymer
Interfacial
Microfluidic
Processing
Wicke–
Kallenbach ~101 ~0 \ 90.0 4.56 12 (13)
γ-alumina
All-vapor
phase
synthesis
Time-lag
~400
~0 Vacuum
381 154 104
This work
~400 214 88 141
~400 368 149 152
Wicke–
Kallenbach
~700
~0
Sweeping,
120-200
mL/min N2
449 157 45
~700 456 164 67
~700 849 305 72
~ 100 621 31 46
~100 880 43 71
~100 1606 79 74
34
Table S2. Feed conditions and separation performance of C3 splitter
Feed pressure 15 bar
Feed vapor fraction 0
Feed flowrate 295.53 mol/s
Feed propylene mol% 70.0 mol%
Column pressure 15
Total stages 247
Feed stage 192
Trays 247
Propylene in distillate 99.7 mol%
Propane in bottoms 99.0 mol%
Reboiler duty 40.6 MW
Condenser duty 40.6 MW
References and Notes 1. D. S. Sholl, R. P. Lively, Seven chemical separations to change the world. Nature 532, 435–
437 (2016). doi:10.1038/532435a Medline 2. K. S. Park, Z. Ni, A. P. Côté, J. Y. Choi, R. Huang, F. J. Uribe-Romo, H. K. Chae, M.
O’Keeffe, O. M. Yaghi, Exceptional chemical and thermal stability of zeolitic imidazolate frameworks. Proc. Natl. Acad. Sci. U.S.A. 103, 10186–10191 (2006). doi:10.1073/pnas.0602439103 Medline
3. H. Furukawa, K. E. Cordova, M. O’Keeffe, O. M. Yaghi, The chemistry and applications of metal-organic frameworks. Science 341, 1230444 (2013). doi:10.1126/science.1230444 Medline
4. H. Bux, F. Liang, Y. Li, J. Cravillon, M. Wiebcke, J. Caro, Zeolitic imidazolate framework membrane with molecular sieving properties by microwave-assisted solvothermal synthesis. J. Am. Chem. Soc. 131, 16000–16001 (2009). doi:10.1021/ja907359t Medline
5. Y. C. Pan, T. Li, G. Lestari, Z. P. Lai, Effective separation of propylene/propane binary mixtures by ZIF-8 membranes. J. Membr. Sci. 390-391, 93–98 (2012). doi:10.1016/j.memsci.2011.11.024
6. H. T. Kwon, H. K. Jeong, A. S. Lee, H. S. An, J. S. Lee, Heteroepitaxially grown zeolitic imidazolate framework membranes with unprecedented propylene/propane separation performances. J. Am. Chem. Soc. 137, 12304–12311 (2015). doi:10.1021/jacs.5b06730 Medline
7. K. Li, D. H. Olson, J. Seidel, T. J. Emge, H. Gong, H. Zeng, J. Li, Zeolitic imidazolate frameworks for kinetic separation of propane and propene. J. Am. Chem. Soc. 131, 10368–10369 (2009). doi:10.1021/ja9039983 Medline
8. R. B. Eldridge, Olefin paraffin separation technology - a review. Ind. Eng. Chem. Res. 32, 2208–2212 (1993). doi:10.1021/ie00022a002
9. M. Galizia, W. S. Chi, Z. P. Smith, T. C. Merkel, R. W. Baker, B. D. Freeman, 50th anniversary perspective: Polymers and mixed matrix membranes for gas and vapor separation: A review and prospective opportunities. Macromolecules 50, 7809–7843 (2017). doi:10.1021/acs.macromol.7b01718
10. M. J. Lee, H. T. Kwon, H. K. Jeong, High-flux zeolitic imidazolate framework membranes for propylene/propane separation by postsynthetic linker exchange. Angew. Chem. Int. Ed. 57, 156–161 (2018). doi:10.1002/anie.201708924 Medline
11. E. Jang, E. Kim, H. Kim, T. Lee, H.-J. Yeom, Y.-W. Kim, J. Choi, Formation of ZIF-8 membranes inside porous supports for improving both their H2/CO2 separation performance and thermal/mechanical stability. J. Membr. Sci. 540, 430–439 (2017). doi:10.1016/j.memsci.2017.06.072
12. M. Drobek, M. Bechelany, C. Vallicari, A. Abou Chaaya, C. Charmette, C. Salvador-Levehang, P. Miele, A. Julbe, An innovative approach for the preparation of confined ZIF-8 membranes by conversion of ZnO ALD layers. J. Membr. Sci. 475, 39–46 (2015). doi:10.1016/j.memsci.2014.10.011
13. A. J. Brown, N. A. Brunelli, K. Eum, F. Rashidi, J. R. Johnson, W. J. Koros, C. W. Jones, S. Nair, Separation membranes. Interfacial microfluidic processing of metal-organic framework hollow fiber membranes. Science 345, 72–75 (2014). doi:10.1126/science.1251181 Medline
14. S. C. Hess, R. N. Grass, W. J. Stark, MOF channels within porous polymer film: Flexible, self-supporting ZIF-8 poly(ether sulfone) composite membrane. Chem. Mater. 28, 7638–7644 (2016). doi:10.1021/acs.chemmater.6b02499
15. T. C. T. Pham, T. H. Nguyen, K. B. Yoon, Gel-free secondary growth of uniformly oriented silica MFI zeolite films and application for xylene separation. Angew. Chem. Int. Ed. 52, 8693–8698 (2013). doi:10.1002/anie.201301766 Medline
16. M. Y. Jeon, D. Kim, P. Kumar, P. S. Lee, N. Rangnekar, P. Bai, M. Shete, B. Elyassi, H. S. Lee, K. Narasimharao, S. N. Basahel, S. Al-Thabaiti, W. Xu, H. J. Cho, E. O. Fetisov, R. Thyagarajan, R. F. DeJaco, W. Fan, K. A. Mkhoyan, J. I. Siepmann, M. Tsapatsis, Ultra-selective high-flux membranes from directly synthesized zeolite nanosheets. Nature 543, 690–694 (2017). doi:10.1038/nature21421 Medline
17. I. Stassen, M. Styles, G. Grenci, H. V. Gorp, W. Vanderlinden, S. D. Feyter, P. Falcaro, D. D. Vos, P. Vereecken, R. Ameloot, Chemical vapour deposition of zeolitic imidazolate framework thin films. Nat. Mater. 15, 304–310 (2016). doi:10.1038/nmat4509 Medline
18. H. T. Kwon, H.-K. Jeong, A. S. Lee, H. S. An, T. Lee, E. Jang, J. S. Lee, J. Choi, Defect-induced ripening of zeolitic-imidazolate framework ZIF-8 and its implication to vapor-phase membrane synthesis. Chem. Commun. (Camb.) 52, 11669–11672 (2016). doi:10.1039/C6CC05433A Medline
19. W. Li, P. Su, Z. Li, Z. Xu, F. Wang, H. Ou, J. Zhang, G. Zhang, E. Zeng, Ultrathin metal-organic framework membrane production by gel-vapour deposition. Nat. Commun. 8, 406 (2017). Medline
20. C. H. Chang, R. Gopalan, Y. S. Lin, A comparative study on thermal and hydrothermal stability of alumina, titania and zirconia membranes. J. Membr. Sci. 91, 27–45 (1994). doi:10.1016/0376-7388(94)00041-7
21. Y. S. Lin, A. J. Burggraaf, Modelling and analysis of CVD processes in porous media for ceramic composite preparation. Chem. Eng. Sci. 46, 3067–3080 (1991). doi:10.1016/0009-2509(91)85010-U
22. M. Tsapatsis, G. R. Gavalas, A kinetic model of membrane formation by CVD of SiO2 and Al2O3. AIChE J. 38, 847–856 (1992). doi:10.1002/aic.690380606
23. Y. C. Pan, W. Liu, Y. J. Zhao, C. Q. Wang, Z. P. Lai, Improved ZIF-8 membrane: Effect of activation procedure and determination of diffusivities of light hydrocarbons. J. Membr. Sci. 493, 88–96 (2015). doi:10.1016/j.memsci.2015.06.019
24. D. F. Liu, X. L. Ma, H. X. Xi, Y. S. Lin, Gas transport properties and propylene/propane separation characteristics of ZIF-8 membranes. J. Membr. Sci. 451, 85–93 (2014). doi:10.1016/j.memsci.2013.09.029
25. N. Hara, M. Yoshimune, H. Negishi, K. Haraya, S. Hara, T. Yamaguchi, Diffusive separation of propylene/propane with ZIF-8 membranes. J. Membr. Sci. 450, 215–223 (2014). doi:10.1016/j.memsci.2013.09.012
26. N. Hara, M. Yoshimune, H. Negishi, K. Haraya, S. Hara, T. Yamaguchi, Effect of temperature on synthesis of ZIF-8 membranes for propylene/propane separation by counter diffusion method. J. Jpn. Petrol. Inst. 58, 237–244 (2015). doi:10.1627/jpi.58.237
27. S. Tanaka, K. Okubo, K. Kida, M. Sugita, T. Takewaki, Grain size control of ZIF-8 membranes by seeding-free aqueous synthesis and their performances in propylene/ propane separation. J. Membr. Sci. 544, 306–311 (2017). doi:10.1016/j.memsci.2017.09.037
28. K. Eum, A. Rownaghi, D. Choi, R. R. Bhave, C. W. Jones, S. Nair, Fluidic processing of high-performance ZIF-8 membranes on polymeric hollow fibers: Mechanistic insights and microstructure control. Adv. Funct. Mater. 26, 5011–5018 (2016). doi:10.1002/adfm.201601550
29. Materials and methods are available as supplementary materials. 30. P. Adhikari, M. Xiong, N. Li, X. Zhao, P. Rulis, W.-Y. Ching, Structure and electronic
properties of a continuous random network model of an amorphous zeolitic imidazolate framework (a-ZIF). J. Phys. Chem. C 120, 15362–15368 (2016). doi:10.1021/acs.jpcc.6b06337
31. H. Tao, T. D. Bennett, Y. Yue, Melt-quenched hybrid glasses from metal-organic frameworks. Adv. Mater. 29, 1601705 (2017). doi:10.1002/adma.201601705 Medline
32. W. J. Koros, R. P. Lively, Water and beyond: Expanding the spectrum of large-scale energy efficient separation processes. AIChE J. 58, 2624–2633 (2012). doi:10.1002/aic.13888
33. J. Yu, Y. C. Pan, C. Q. Wang, Z. P. Lai, ZIF-8 membranes with improved reproducibility fabricated from sputter-coated ZnO/alumina supports. Chem. Eng. Sci. 141, 119–124 (2016). doi:10.1016/j.ces.2015.10.035
34. N. Hara, M. Yoshimune, H. Negishi, K. Haraya, S. Hara, T. Yamaguchi, Thickness reduction of the zeolitic imidazolate framework-8 membrane by controlling the reaction rate during the membrane preparation. J. Chem. Eng. of Jpn 47, 770–776 (2014). doi:10.1252/jcej.14we340
35. N. Hara, M. Yoshimune, H. Negishi, K. Haraya, S. Hara, T. Yamaguchi, ZIF-8 membranes prepared at miscible and immiscible liquid-liquid interfaces. Microporous Mesoporous Mater. 206, 75–80 (2015). doi:10.1016/j.micromeso.2014.12.018
36. N. Hara, M. Yoshimune, H. Negishi, K. Haraya, S. Hara, T. Yamaguchi, Effect of solution concentration on structure and permeation properties of ZIF-8 membranes for propylene/propane separation. J. Chem. Eng. of Jpn 49, 97–103 (2016). doi:10.1252/jcej.15we038
37. H. T. Kwon, H. K. Jeong, In situ synthesis of thin zeolitic-imidazolate framework ZIF-8 membranes exhibiting exceptionally high propylene/propane separation. J. Am. Chem. Soc. 135, 10763–10768 (2013). doi:10.1021/ja403849c Medline
38. H. T. Kwon, H. K. Jeong, Highly propylene-selective supported zeolite-imidazolate framework (ZIF-8) membranes synthesized by rapid microwave-assisted seeding and secondary growth. Chem. Commun. (Camb.) 49, 3854–3856 (2013). doi:10.1039/c3cc41039k Medline
39. H. T. Kwon, H. K. Jeong, Improving propylene/propane separation performance of zeolitic-imidazolate framework ZIF-8 Membranes. Chem. Eng. Sci. 124, 20–26 (2015). doi:10.1016/j.ces.2014.06.021
40. K. Eum, C. Ma, A. Rownaghi, C. W. Jones, S. Nair, ZIF-8 membranes via interfacial microfluidic processing in polymeric hollow fibers: Efficient propylene separation at elevated pressures. ACS Appl. Mater. Interfaces 8, 25337–25342 (2016). doi:10.1021/acsami.6b08801 Medline
41. K. V. Agrawal, B. Topuz, Z. Jiang, K. Nguenkam, B. Elyassi, L. F. Francis, M. Tsapatsis, M. Navarro, Solution-processable exfoliated zeolite nanosheets purified by density gradient centrifugation. AIChE J. 59, 3458–3467 (2013). doi:10.1002/aic.14099
42. V. Gokhale, S. Hurowitz, J. B. Riggs, A comparison of advanced distillation control techniques for a propylene/propane splitter. Ind. Eng. Chem. Res. 34, 4413–4419 (1995). doi:10.1021/ie00039a033
43. U. Lee, J. Kim, I. Seok Chae, C. Han, Techno-economic feasibility study of membrane based propane/propylene separation process. Chem. Eng. Process. Process Intensif. 119, 62–72 (2017). doi:10.1016/j.cep.2017.05.013
44. T. M. Zygula, K. Kolmetz, Design considerations for propylene splitters, AIChE Spring Meeting and Global Congress on Process Safety, Chicago, 16 March 2011.
45. S. S. Jogwar, P. Daoutidis, Dynamics and control of vapor recompression distillation. J. Process Contr. 19, 1737–1750 (2009). doi:10.1016/j.jprocont.2009.07.001
46. D. F. Schneider, Heat integration complicates heat pump troubleshooting. Hydrocarbon Process. 81, 53–56 (2002).
47. ICIS, Americas Chemicals Outlook 2017 (ICIS, 2017). 48. Y. Huang, T. C. Merkel, R. W. Baker, Pressure ratio and its impact on membrane gas
separation processes. J. Membr. Sci. 463, 33–40 (2014). doi:10.1016/j.memsci.2014.03.016
49. R. W. Baker, Membrane Technology and Applications (John Wiley & Sons, ed. 3, 2012). 50. R. W. Baker, A. R. Da Costa, R. Daniels, I. Pinnau, Z. He, Membrane-augmented
polypropylene manufacturing, U.S. patent 6,271,319 B1 (2000). 51. M. T. Castoldi, J. C. Pinto, P. A. Melo, Modeling of the separation of propene/propane
mixtures by permeation through membranes in a polymerization system. Ind. Eng. Chem. Res. 46, 1259–1269 (2007). doi:10.1021/ie060333q
52. Global Polypropylene (PP) Market report 2017: $100+ Billion market size, demand forecasts, industry trends and updates 2016–2022—Research and markets (2017).
53. M. S. Peters, K. D. Timmerhaus, R. E. West, Plant Design and Economics for Chemical Engineers (McGraw-Hill, 2003).
54. I. K. Kookos, Optimal design of membrane/distillation column hybrid processes. Ind. Eng. Chem. Res. 42, 1731–1738 (2003). doi:10.1021/ie020616s
55. Y. Naidu, R. K. Malik, A generalized methodology for optimal configurations of hybrid distillation-pervaporation processes. Chem. Eng. Res. Des. 89, 1348–1361 (2011). doi:10.1016/j.cherd.2011.02.025
56. H. J. Salgado-Gordon, G. Valbuena-Moreno, Technical and economic evaluation of the separation of light olefins (ethylene and propylene) by using π-complexation with silver salts. Ciencia Tecnol. Y Futur. 4, 73–87 (2011).
57. A. A. Kiss, S. J. Flores Landaeta, C. A. Infante Ferreira, Towards energy efficient distillation technologies - Making the right choice. Energy 47, 531–542 (2012). doi:10.1016/j.energy.2012.09.038
58. J. R. Alcántara-Avila, F. I. Gomez-Castro, J. G. Segovia-Hernandez, K. Sotowa, T. Horikawa, Optimal design of cryogenic distillation columns with side heat pumps for the propylene/propane separation. Chem. Eng. Process. Process Intensif. 82, 112–122 (2014). doi:10.1016/j.cep.2014.06.006
59. Z. Olujić, L. Sun, A. de Rijke, P. J. Jansens, Conceptual design of an internally heat integrated propylene-propane splitter. Energy 31, 3083–3096 (2006). doi:10.1016/j.energy.2006.03.030
60. A. A. Shenvi, D. M. Herron, R. Agrawal, Energy efficiency limitations of the conventional heat integrated distillation column (HIDiC) configuration for binary distillation. Ind. Eng. Chem. Res. 50, 119–130 (2011). doi:10.1021/ie101698f
61. J. Park, K. Kim, J. W. Shin, K. Tak, Y. K. Park, Performance study of multistage membrane and hybrid distillation processes for propylene/propane separation. Can. J. Chem. Eng. 95, 2390–2397 (2017). doi:10.1002/cjce.22914
62. D. E. Gottschlich, D. L. Roberts, “Energy minimization of separation processes using conventional/membrane hybrid systems,” no. DOE/ID-10301 (1990); www.osti.gov/servlets/purl/6195331.
63. W. Stephan, R. D. Noble, C. A. Koval, Design methodology for a membrane/distillation column hybrid process. J. Membr. Sci. 99, 259–272 (1995). doi:10.1016/0376-7388(94)00255-W
64. T. Pettersen, K. M. Lien, Design of hybrid distillation and vapor permeation processes. J. Membr. Sci. 102, 21–30 (1995). doi:10.1016/0376-7388(94)00216-L
65. T. Pettersen, A. Argo, R. D. Noble, C. A. Koval, Design of combined membrane and distillation processes. Sea Technol. 6, 175–187 (1996). doi:10.1016/0956-9618(96)00151-8
66. T. G. Pressly, K. M. Ng, A break-even analysis of distillation-membrane hybrids. AIChE J. 44, 93–105 (1998). doi:10.1002/aic.690440111
67. J. A. Caballero, I. E. Grossmann, M. Keyvani, E. S. Lenz, Design of hybrid distillation-vapor membrane separation systems. Ind. Eng. Chem. Res. 48, 9151–9162 (2009). doi:10.1021/ie900499y
68. S. Pedram, T. Kaghazchi, M. T. Ravanchi, Performance and energy consumption of membrane-distillation hybrid systems for olefin-paraffin separation. Chem. Eng. Technol. 37, 587–596 (2014). doi:10.1002/ceat.201200621
69. J. Ploegmakers, A. R. T. Jelsma, A. G. J. van der Ham, K. Nijmeijer, Economic evaluation of membrane potential for ethylene/ethane separation in a retrofitted hybrid membrane-distillation plant using unisim design. Ind. Eng. Chem. Res. 52, 6524–6539 (2013). doi:10.1021/ie400737s
70. A. Motelica, O. S. L. Bruinsma, R. Kreiter, M. den Exter, J. F. Vente, Membrane retrofit option for paraffin/olefin separation-a technoeconomic evaluation. Ind. Eng. Chem. Res. 51, 6977–6986 (2012). doi:10.1021/ie300587u
71. D. K. Kim, S. B. Lee, P. Yoon, Numerical simulation of fixed-bed catalytic reactor for isopropyl alcohol synthesis. Korean J. Chem. Eng. 6, 99–104 (1989). doi:10.1007/BF02697486
72. Y. Xu, K. T. Chuang, A. R. Sanger, Design of a process for production of isopropyl alcohol by hydration of propylene in a catalytic distillation column. Chem. Eng. Res. Des. 80, 686–694 (2002). doi:10.1205/026387602760312908
73. A. Shamiri, M. A. Hussain, F. S. Mjalli, M. S. Shafeeyan, N. Mostoufi, Experimental and modeling analysis of propylene polymerization in a pilot-scale fluidized bed reactor. Ind. Eng. Chem. Res. 53, 8694–8705 (2014). doi:10.1021/ie501155h
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