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

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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

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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

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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

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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

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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

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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)

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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

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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

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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

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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

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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)

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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.

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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.

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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.

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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.

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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.

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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

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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.

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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.

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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

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