Journal of Membrane Science 432 (2013) 106–114

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Journal of Membrane Science

0376-73

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n Corr

E-m

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Dehydration of natural gas using membranes. Part II: Sweep/countercurrentdesign and field test

Haiqing Lin n, Scott M. Thompson, Adrian Serbanescu-Martin, Johannes G. Wijmans, Karl D. Amo,Kaaeid A. Lokhandwala, Bee Ting Low, Timothy C. Merkel

Membrane Technology and Research, Inc., 39630 Eureka Drive, Newark, CA 94560, USA

a r t i c l e i n f o

Article history:

Received 20 October 2012

Received in revised form

18 December 2012

Accepted 28 December 2012Available online 9 January 2013

Keywords:

Membrane

Natural gas dehydration

Sweep/countercurrent spiral-wound

module

Process designs

88/$ - see front matter & 2013 Elsevier B.V. A

x.doi.org/10.1016/j.memsci.2012.12.049

esponding author. Tel.: þ1 650 543 3359; fax

ail address: haiqing.lin@mtrinc.com (H. Lin).

a b s t r a c t

Membrane technology is an attractive alternative to conventional glycol dehydrators for natural gas

dehydration. While membranes with high water/methane selectivity are available, the separation

performance of membrane systems is restricted by the moderate feed-to-permeate pressure ratio

encountered in this application, resulting in high membrane area requirement and high methane loss.

This paper (Part II of a two-part study) addresses that challenge by evaluating suitable process designs

to overcome the limitations of feed-to-permeate pressure ratio, using a process simulation package. A

design with a dry gas on the permeate side as a countercurrent sweep shows the best combination of

low methane loss and minimum membrane area requirement. The design was experimentally

evaluated by field testing a spiral-wound module that incorporates countercurrent operation, at a

natural gas processing plant. The use of a dry sweep gas on the permeate side significantly increased

water vapor flux across the membrane. For example, as sweep/feed flow rate ratio increased from 0% to

5.9% at a permeate pressure of 4.8 bar, water flux increased by 2.5 times, from 0.13 slpm/m2 (standard

liter per minute per m2 membrane area) to 0.47 slpm/m2. The measured water flux reached about 80%

of the value calculated with the process simulation package for a perfect sweep/countercurrent module,

confirming the suitability of the spiral-wound modules with countercurrent configuration for natural

gas dehydration.

& 2013 Elsevier B.V. All rights reserved.

1. Introduction

Membrane technology is an attractive alternative to conven-tional glycol dehydrators for natural gas dehydration, due to theinherent advantages of membranes such as small footprint,passive operation without moving parts, and good reliability foroperation in remote locations [1–5]. Part I of this study describedthin film composite membranes based on microphase-separatedblock copolymers (Pebaxs) with promising H2O/CH4 separationproperties [6]. The composite membranes were extensively testedin the laboratory by varying operating parameters including feedgas composition, feed gas flow rate and sweep gas flow rate. Themicroporous support and paper layers of the composite mem-branes contributed the main transport resistance to water vapor,which was reduced by using a dry gas as sweep on the permeateside. At high sweep flow rates, the membrane showed watervapor permeance up to 2000 gpu [1 gpu¼10�6 cm3(STP)/cm2 scm Hg] and H2O/CH4 selectivity up to 1500 [6].

ll rights reserved.

: þ1 650 328 6580.

The chemical structure of Pebaxs copolymers, shown in Fig. 1,is micro-phase separated and contains bi-continuous phases: ahydrophilic polyether phase and a rigid glassy polyamide phase[7–9]. The hydrophilic polyether phase provides good H2O/CH4

separation properties, and the rigid glassy phase provides goodmechanical properties for operation at high pressure and highwater content [7–12].

Part II of this study describes the efforts to bring the mem-brane technology from laboratory-scale study to commercializa-tion by performing membrane process design studies andconducting a field test at a natural gas processing plant. Severalmembrane process designs were evaluated, and commercialsuitability of the most promising one was further verified by afield test at a natural gas processing plant.

2. Membrane process designs for natural gas dehydration

Successful commercialization of any membrane technologyrelies on an optimal process design to reduce capital and operat-ing costs and achieve high recovery of products [13–18]. Fig. 2shows four typical membrane process designs considered fornatural gas dehydration. The feed gas contains 1000 ppm water

C PA C O PE O

O O

Hn

HO C PA C O PE O

O O

Hn

HO

Fig. 1. Structure of a Pebaxs block copolymer. PA is a polyamide segment such as

Nylon 6 in Pebaxs 1657 and Nylon 12 in Pebaxs 2533; PE is a polyether segment

such as poly(ethylene oxide) in Pebaxs 1657 and poly(tetramethylene oxide) in

Pebaxs 2533 [7,8].

Fig. 2. Four membrane processes for dehydration of a binary H2O/CH4 mix-

ture from 1000 ppm to 100 ppm with potential commercial application.

(a) Conventional cross-flow design with the permeate stream at atmospheric

pressure (15% CH4 loss); (b) cross-flow design with the permeate stream under

vacuum (1.9% CH4 loss); (c) countercurrent design using an expanded slipstream

from the residue as sweep (7.0% CH4 loss); and (d) countercurrent design using

a dry gas stream as sweep (0.78% CH4 loss). The feed gas has a flow rate of

1.5 m3(STP)/s at 68.7 bar.

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114 107

in methane at 68.7 bar and has a flow rate of 1.5 m3(STP)/s. Themembrane processes reduce water content to 100 ppm, which isthe typical requirement of natural gas pipeline specifications.A commercial process simulator, ChemCAD 6.4 (ChemStations,Houston, TX), used in conjunction with a membrane unit opera-tion add-in, was used to simulate the process conditions. Themembrane used has a water vapor permeance of 1000 gpu andwater/methane selectivity of 500. The simulation results are alsoshown in Fig. 2 and discussed below.

2.1. Conventional cross-flow design

Fig. 2(a) shows a conventional membrane process design with across-flow operating mode that reduces water content from1000 ppm to 100 ppm. The permeate stream can be vented or usedas fuel. Most industrial gas separation membranes operate in cross-flow mode, where the gas on the permeate side of the membrane iscompletely mixed and gas composition is the same for the entiremodule element, even though the gas composition in the upstreamfeed side changes from the feed entrance to the residue exit. Asshown in Fig. 2(a), this design leads to a methane loss of 15%, whichis much higher than the loss in conventional glycol dehydrators(1%). The high methane loss is caused mainly by the nature of theseparation, which is limited by the moderate pressure ratio of thefeed to permeate streams, regardless of the membrane performance[1,3]. The importance of pressure ratio for natural gas dehydration isexplained as follows [1]. Considering a separation with a feedpressure of p2 and permeate pressure of p1, a flow of water vaporacross the membrane can only occur if the water partial pressure onthe feed side of the membrane is greater than that on the permeateside of the membrane:

x2,H2Op24x1,H2Op1 ð1Þ

where x1,H2O and x2,H2O are the water molar concentration in thepermeate and feed, respectively. The maximum water concentrationenrichment achieved by the membrane can be expressed as

x1,H2O

x2,H2Oo

p2

p1ð2Þ

This means that the enrichment of water can never exceed thepressure ratio of p2=p1, no matter how selective the membraneis [1]. The balance of the permeate stream is methane and higherhydrocarbons.

As shown in Fig. 2(a), the average feed-to-residue water contentis about 391 ppm (a logarithmic mean average of 1000 ppm in thefeed and 100 ppm in the residue) [6]. With a feed-to-permeatepressure ratio of 15, the maximum permeate water content wouldbe about 6000 ppm, which is very close to the simulated result. Thepermeate stream can be compressed to 68.7 bar and, after watercondensation, recycled to the feed of the membrane to reduce themethane loss. However, the capital and operating costs of thiscompression diminish the advantages of the membrane system,and make it prohibitively expensive in comparison to glycoldehydrators [1].

2.2. Cross-flow design with the permeate under vacuum

One simple way to overcome the pressure-ratio limit onnatural gas dehydration is to decrease the permeate pressureusing a vacuum pump, as shown in Fig. 2(b) [4,5,19]. By decreas-ing the permeate pressure from 4.5 bar to 0.45 bar, the feed-to-permeate pressure ratio increases from 15 to 150, resulting in asignificant decrease in the required membrane area and asso-ciated methane loss. The design of Fig. 2(b) has a membrane areaof 280 m2 and a methane loss of 1.9%, which is a significantimprovement compared with the conventional design inFig. 2(a) that requires a membrane area of 2400 m2 and resultsin a methane loss of 15%. However, the requirement of a vacuumpump in design 2(b) increases the capital and operating cost of amembrane system. More importantly, a leak of air into thevacuum will introduce oxygen into the system, which needs tobe closely monitored and mitigated for safety reasons [4].

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114108

2.3. Countercurrent design using an expanded residue slipstream as

sweep

Another way to mitigate the influence of the pressure-ratiolimit on the separation is to use a dry sweep gas in the permeate,which flows countercurrently to the feed gas [2,17,20–23]. InFig. 2(c), a slipstream of the dehydrated residue gas is expandedto 4.5 bar and used as a permeate sweep gas. As it flows inthe permeate flow channel, the sweep gas mixes with the waterand methane permeating from the high pressure feed stream.The outlet stream on the permeate side is the combination ofthe slipstream of the dehydrated residue gas, plus the methaneand water that permeate from the feed (and is called the‘‘sweepþpermeate’’ stream). The simulated methane loss is7.0%, which is lower than the conventional cross-flow designshown in Fig. 2(a). However, the methane loss of 7.0% is still muchhigher than that for conventional glycol dehydrators (withmethane loss of 1%). This process design can only be adopted ifthe methane in the permeate can be fully utilized as fuel atthe site.

2.4. Countercurrent design using dry nitrogen as sweep

Instead of using a slipstream from the dried residue gas assweep [shown in Fig. 2(c)], a dry gas stream such as nitrogen canbe used as the permeate sweep, as shown in Fig. 2(d). Many gasprocessing plants have dry venting or fuel gas available for use asthe sweep. For example, there may be dry nitrogen gas producedon-site for other applications such as blanketing hydrocarbonstorage tanks. This dry nitrogen gas can be used to sweep thepermeate in the dehydration modules before being used toblanket the tanks. The design assumes a permeate pressure of4.5 bar.

The use of a dry sweep gas decreases both the requiredmembrane area and the methane loss. This design has thecombined advantages of those shown in Fig. 2(b) and (c); itavoids the use of the vacuum pump while providing low methaneloss. For the membrane with 10% sweep/feed flow rate ratio, themethane loss is only 0.78% and the required membrane area isonly 120 m2.

In this design, the separation is not limited by the feed-to-permeate pressure ratio, as indicated by the much lower watervapor partial pressure in the ‘‘sweepþpermeate’’ stream(0.0374 bar) compared to that in the feed (0.0687 bar). Instead,the water vapor in the ‘‘sweepþpermeate’’ stream is limited by

100

1000

100 1000

Req

uire

d m

embr

ane

area

(m2 )

2b. vacuum permeate

2d. dry N2sweep

2a. conventional crossflow

2c. dry residue gas sweep

5050

H2O/CH4 selectivity

Fig. 3. Comparison of (a) the membrane area required and (b) methane loss in the dehy

permeate operation [Fig. 2(b)], dry residue gas sweep operation [Fig. 2(c)], and dry N

product gas contains 100 ppm water.

the sweep-to-feed flow rate ratio. The water vapor flow rate in the‘‘sweepþpermeate’’ stream must be less than the feed watervapor flow rate. Increasing sweep gas flow rate can reduce themembrane area and methane loss.

2.5. Evaluation of four membrane system designs

Fig. 3 compares the membrane system performance (includingthe required membrane area and methane loss) of the fourproposed membrane dehydration processes as a function ofH2O/CH4 selectivity. The H2O/CH4 selectivity is varied by chan-ging CH4 permeance and keeping H2O permeance constant at1000 gpu. The detailed operating conditions are shown in Fig. 2.In general, higher H2O/CH4 selectivity indicates lower CH4 per-meance, resulting in an increase in the required membrane areabut a reduction in the CH4 loss. This can be explained as below.Higher H2O/CH4 selectivity leads to higher water enrichment orhigher water permeate concentration. As water permeate con-centration increases, the driving force for water vapor permeationacross the membrane decreases and therefore, the requiredmembrane area to remove water vapor from 1000 ppm in thefeed to 100 ppm in the residue increases. On the other hand, theCH4 permeate flow rate (or CH4 loss) decreases, probably becausethe decrease in CH4 permeance is more dominant than theincrease in the membrane area.

The membrane area determines the quantity of membranemodules, module housings, piping and instrumentationneeded [1]. Therefore, the capital cost of a membrane systemoften scales linearly with the membrane area requirement. Evenfor the cross-flow design with permeate stream under vacuum[see Fig. 2(b)], the cost of the vacuum pump (at a powerrequirement of 10 kW) is a small portion of the overall membranesystem cost (with a membrane area of 280 m2). General rules ofthumb for membrane system cost estimates are $500/m2 mem-brane area for a high-pressure membrane skid [1], and $1000/kWfor a vacuum pump. Therefore, increasing H2O/CH4 selectivityincreases the capital cost of the membrane system for the fourprocess designs shown in Fig. 2, while the countercurrent designwith a dry stream as sweep [Fig. 2(d)] shows only a marginalincrease in the membrane area required as H2O/CH4 selectivityincreases.

The countercurrent design with dry gas as sweep [Fig. 2(d)]shows the best combination of low CH4 loss and minimalmembrane area requirement, among the four process designsconsidered. For membranes with H2O/CH4 selectivity higher than

0.1

1

10

100 1000

Met

hane

loss

(%)

2b. vacuum permeate

2d. dry N2sweep

2a. conventional crossflow

2c. dry residue gas sweep

50H2O/CH4 selectivity

dration process designs with conventional cross-flow operation [Fig. 2(a)], vacuum

2 sweep operation [Fig. 2(d)]. Feed gas contains 1000 ppm water at 68.7 bar and

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114 109

400, the process using dry N2 sweep [Fig. 2(d)] shows CH4 lossless than 1%, which is comparable to that of conventional glycolabsorption (with a CH4 loss of 1%). The methane loss can befurther reduced by increasing the sweep/feed flow rate ratio. Forexample, by increasing the sweep/feed flow rate ratio from 10% to20% at a H2O/CH4 selectivity of 500, the methane loss declinesfrom 0.78% to 0.58%, and the required membrane area is reducedfrom 120 m2 to 90 m2. However, the key to success of thecountercurrent process design is the availability of the dry inertgas as sweep. If the dry inert gas needs to be generated fordehydration application only (such as by nitrogen generationusing membrane systems), the cost of the dry gas production canbe prohibitive.

Despite industrial interest in sweep/countercurrent operation[2,17,20–23], the efficiency of countercurrent flow in modules isnot well understood [24]. For example, it has been argued thatcountercurrent design does not offer any better separation effi-ciency than cross-flow, because the porous support layer preventsmixing of the local permeate fluxes (also known as internalconcentration polarization) [25]. Internal concentration polariza-tion has been observed in other gas separations [26,27] and inforward osmosis [28]. More importantly, spiral-wound modulesmay not provide a perfect countercurrent flow pattern [29–32]. Inthis study, a field test of sweep/countercurrent spiral-woundmodules for natural gas dehydration was performed to determinethe impact of this issue.

3. Experimental

3.1. Membrane and module preparation

The thin film composite membranes prepared in this workwere comprised of a paper support layer, a microporous supportlayer, a gutter layer and a thin Pebaxs 2533 selective layer. Themicroporous support layer was formed by phase inversion using acasting machine [3]. The gutter layer and selective layer wereformed by a dip-coating process using an industrial-scalecoater [3].

Fig. 4 shows an exploded view of a conventional cross-flowspiral-wound module for gas separation with three ports (feed,residue and permeate) [3]. Two flat sheets with permeate spacersin between are glued along three sides to form an envelope with

Fig. 4. Exploded view of a conventional spiral-wound membrane module (in

cross-flow mode) for gas separation [3,33].

an open end facing the permeate pipe. Many of these envelopesare wound around a perforated central collection pipe with a feedspacer between each set of envelopes. Feed gas passes down themodule parallel to the permeate pipe in the channels created bythe feed spacer. A portion of the feed gas permeates the mem-brane, enters the permeate channel, and exits through thecollection pipe.

A new, modified four-port sweep/countercurrent moduleallows for the introduction of sweep gas into the permeatechannel. As shown in Fig. 5, only two modifications are requiredto convert a conventional spiral-wound module (which operatesin cross-flow) to a sweep/countercurrent module [29,32,34,35].First, the permeate pipe is plugged in the middle, forming twoseparate compartments. Second, during module rolling, oneadditional glue line is applied in the middle of the permeatechannel. This glue line extends from the permeate pipe to themiddle of the membrane envelope. As shown in Fig. 5(b), thesemodifications allow the permeate channel to be swept with asweep gas and the module to operate in a partial countercurrentmode [29]. Fig. 5(b) also shows that the module does not have aperfect countercurrent flow pattern between the feed and perme-ate flow. Nevertheless, it will be shown later in this study that themodule can still have very high sweep efficiency, despite thepresence of a non-ideal countercurrent flow pattern.

During the field test conducted for this study, one sweep/countercurrent module containing 3.0 m2 membrane area wastested; it was 4 in. in diameter and 40 in. long. This module wasconfirmed to be defect-free by measuring pure-gas CO2/CH4

selectivity, which is shown in detail in Section 4.2 (Pure- andmixed-gas permeance in the modules). The effects of the counter-current flow pattern and the operating conditions (such as feedand permeate pressure, and feed and sweep flow rates) onmodule performance were evaluated.

3.2. Membrane pilot-scale test skid

A pilot-scale membrane skid was used to test various modulesfor water removal from natural gas. A photograph of the skid isshown in Fig. 6(a). Fig. 6(b) shows the simplified process flowdiagram. In the test, feed gas entering the skid passed through acoalescing filter to remove entrained liquids and particulateslarger than 5 mm. The gas was then routed to a pressure vesselaccommodating a 4-inch-diameter sweep/countercurrent spiral-wound module. The feed flow rates were controlled by flowcontrol valves located at the entrance to the tube vessel. Byvarying the feed flow rates and feed and permeate pressures,water removal rates could be varied in the range of 10% to 90þ%.

The system treated natural gas flowing at rates between 0.03and 0.2 m3(STP)/s at feed pressures up to 60 bar. The temperature,pressure, and flow rate of the feed, residue, and permeate gasstreams were measured using the gauges on the skid. Gascompositions were determined with an Agilent MicroGC 3000portable gas chromatograph, which detected hydrocarbons C1–C9,CO2 and nitrogen that were present, and measured their concen-trations to levels as low as 10 ppmv. Water vapor concentrationwas determined using Draeger tubes (Draeger Safety Inc., Pitts-burgh, PA), an industry standard measurement method.

3.3. Field test site information

The pilot-scale membrane skid was installed at a natural gasprocessing plant in Isleton, CA, owned by Towne Exploration Com-pany. The plant gathered natural gas from several gas wells nearby.Details of the gas streams treated are given in Table 1. A typical gasstream entering the processing plant at 30–60 bar was saturated withwater (600–1200 ppm), and contained 17% nitrogen, 79% methane,

Fig. 5. View of the membrane envelope for two types of spiral-wound modules. The flow pattern in the conventional module (a) is cross-flow, whereas the modified

module (b) accepts a sweep gas on the permeate side and operates in a partial countercurrent pattern [29,32,34].

Fig. 6. (a) Photograph and (b) process flow diagram of the natural gas dehydration field test system at a natural gas processing plant. The system includes two 4-inch

modules installed vertically. One module housing was modified to accommodate the sweep/countercurrent module.

Table 1Typical conditions and compositions of gas streams used at Towne Exploration’s

Isleton site. The stream compositions were measured using a micro-gas

chromatograph.

Operating conditions Inlet gas Sweep gas

Flow rate [m3(STP)/s] 0.05–0.15 0–0.10

Pressure (bar) 30–60 0–60

Temperature (1C) 25–40 25–40

Water content (ppm) 600–1200 o10

Gas compositions (mol%)Nitrogen 16.86 48.85

Carbon dioxide 0.31 0.05

Methane 79.23 50.31

Ethane 1.70 0.30

Propane 0.94 0.14

i-Butane 0.20 0.03

n-Butane 0.24 0.04

i-Pentane 0.10 0.02

n-Pentane 0.08 0.03

Hexaneþ 0.22 0.22

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114110

1.7% ethane, and 0.9% propane, with the balance (1.4%) being higherhydrocarbons. The sweep gas was a vent gas from a membranesystem for nitrogen removal from natural gas and had an undetect-able level of water content [36]. After passing the feed through a filterseparator to remove any entrained liquids and particulates, a slip-stream of the raw natural gas [about 0.05 m3(STP)/s] entered themembrane test skid. After the feed gas passed through the module,the residue gas flowed through a back pressure regulator and wasused as fuel for plant compressors at 18 bar. The permeate gas wasvented at the site. However, in the larger-scale practical operation, thepermeate gas often cannot be vented. It should be considered as fuelor sent to the flare line.

3.4. Determination of gas permeance

Pure-gas flux through composite membranes is oftendescribed by permeance, (P/‘)A, which is given by [3]:

P

‘

� �A

�NA

Am p2�p1Þ� ð3Þ

where NA is the steady state flow rate of gas A through themembrane, ‘ is the thickness of the membrane selective layer, andAm is the membrane area. Permeance is often expressed in gaspermeation units (gpu), where 1 gpu¼10�6 cm3(STP)/(cm2 s cm Hg).

The ideal selectivity of a membrane for gas A over gas B is theratio of the pure-gas permeances:

aA=B ¼P=‘� �

A

P=‘� �

B

ð4Þ

The mixed-gas permeance of component A in the module,P=‘� �

A, can be calculated using the following equation for the

cross-flow module operation, assuming the permeate in themodule is well mixed:

P

‘

� �A

�NPx1,A

Am p2x2,A�p1x1,A Þ� ð5Þ

where NP is the total permeate flow rate, x1,A is the mole fraction ofcomponent A in the permeate and x2,A is the average concentrationof component A in the membrane upstream. While the parametersof NP, x1,A, Am, p2 and p1 can be measured directly using theinstruments on the membrane skid, the parameter of x2,A needsto be estimated as the logarithmic mean of the component A

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114 111

concentration on the membrane upstream.

x2,A ¼xF,A�xR,A

ln xF,A=xR,A

� � ð6Þ

where xF,A and xR,A are the mole fraction of component A in thefeed entering the module and the residue exiting the module,respectively.

For the module operating in countercurrent mode, the follow-ing equation can be used to estimate the component permeances[6]:

P

‘

� �A

¼NPx1,AUln DpA,L=DpA,0

� �Am DpA,L�DpA,0

� � ð7Þ

Eq. (7) is similar to Eq. (5), except that the partial pressuredifference is expressed as the logarithmic mean of DpA,L and DpA,0.DpA,0 is the partial pressure difference of component A betweenthe feed inlet and sweepþpermeate stream, as shown in Fig. 2(d).DpA,L is the partial pressure difference of component A betweenthe residue and sweep stream.

Both Eqs. (5) and (7) are derived based on assumptions thatthe flow rates on the upstream and downstream sides areconstant along the flow direction. In the field test, high feed flowrates were used, and therefore, the residue flow rates were oftenabout 95% that of the feed gas. Therefore, both equations can beused to calculate the mixed-gas permeance. The permeancevalues obtained were also confirmed using the ChemCAD processsimulator, which considers the flow rate changes in the feed andpermeate channels.

Table 2Gas permeance and selectivity of the membrane module during the field test at a

natural gas processing plant. The feed gas had a flow rate of 0.050 m3(STP)/s at

35 bar and 27 1C with the gas composition shown in Table 1. The permeate

pressure was near atmospheric pressure. The stage-cut was 2.5%.

Gas component Permeance (gpu) Gas/CH4 selectivity

N2 12 0.62

CH4 19 –

CO2 57 2.9

C2H6 38 2.0

C3H8 58 3.0

i-C4H10 62 3.3

n-C4H10 100 5.3

i-C5H12 110 5.8

n-C5H12 140 7.4

Hexaneþ 95 4.9

4. Field test results

4.1. Membrane production scale-up and water vapor permeance

The production of thin film composite membranes based onPebaxs polymers was scaled up to demonstrate a successful com-mercial scale production run using commercial coating machines. Themembrane configuration and coating conditions were optimized toachieve good production reproducibility.

The resulting composite membranes were tested for water vaporpermeation at 40 1C and feed relative humidity of 40% using anapparatus similar to that described in the Part I study [6]. As anexample, one batch of membrane showed pure-gas CH4 permeance of20 gpu, water vapor permeance of 930 gpu and water/CH4 selectivityof 47. The water/CH4 selectivity of these scaled-up membranes wasmuch lower than that of membranes prepared in the laboratory,which showed water/CH4 selectivity as high as 1500, as reported inthe Part I study [6]. The discrepancy in the production scale upperformance is due to the use of a gutter layer, which is hydrophobicand shows significant resistance to water vapor permeation. Never-theless, the focus of this study was to investigate the effect ofcountercurrent sweep on the water vapor removal. The membraneconfiguration will be further optimized in future studies, to minimizethe water transport resistance in the support layers.

4.2. Pure- and mixed-gas permeance in the modules

The sweep module was tested in the laboratory for quality controlpurposes, before the field test began. Pure-gas permeances weremeasured at 22 1C and feed pressures of about 10 psig. The moduleshowed pure-gas CO2/CH4 selectivity of 5.9, indicating that themodule was defect-free. For an initial integrity test, the module wasexposed to compressed air at 900 psig for 2 h. The pure-gas per-meance and selectivity did not change after the high pressureexposure, indicating the stability of the module at high pressures.

Mixed-gas permeance and selectivity of these modules duringfield tests can be used as an indicator for module integrity in a realnatural gas environment. Table 2 shows representative data for thesweep/countercurrent module. The data were taken after the testsreached steady state. Permeances were calculated based on themeasured flow rates, pressures, temperatures, and compositions ofeach stream, using Eqs. (5) and (7). Permeance values were alsopositively confirmed using the ChemCAD process simulator.

In general, mixed-gas methane permeance was higher than thepure-gas methane permeance, which is 7.5 gpu measured at 10 psigand 20 1C. This was expected since the presence of higher hydro-carbons such as butane can significantly plasticize the membrane,leading to an increase in methane permeance. Nevertheless, hydro-carbon/methane selectivities were more or less consistent from testto test. For example, ethane/methane selectivity was about 2.0–2.4for all the tests. These results were consistent with the laboratoryresults for the membrane stamps with the same configuration. Thefield test membrane module performed well and stably in theobserved range of pressures and stage-cuts. In broader terms, thefield tests demonstrated the robustness of the membrane modules inrich natural gas environments (high pressure and high C3þ content).

4.3. Water removal using a sweep module

The sweep module was tested using a dry sweep gas on thepermeate side of the membrane. The feed gas contained 1100 ppmwater vapor and had a flow rate of 0.050 m3(STP)/s at 35 bar. Thesweep gas was a vent gas available at the plant that containednitrogen, light hydrocarbons and essentially no water (that is, lessthan 10 ppm, the detection limit for Draeger tubes). The flow rate andpressure of the sweep gas were varied to investigate their effect onwater removal efficiency.

Fig. 7 shows the effect of permeate pressure on water flux withand without sweep gas. Decreasing the permeate pressure increasedthe water flux. As permeate pressure decreased from 7.9 bar to1.1 bar, water flux increased from 0.069 slpm/m2 to 0.31 slpm/m2,an increase of 340%. The use of sweep flow substantially increasedwater flux (top curve versus bottom curve in Fig. 7). For example, at apermeate pressure of 7.9 bar, a sweep stream flowing at 10% of thefeed flow rate (sweep/feed flow rate ratio of 0.1) increased water fluxfrom 0.069 slpm/m2 to 0.41 slpm/m2.

Fig. 8(a) compares the measured water vapor flux through thefield test module at permeate pressures of 4.8 bar and 7.9 bar as afunction of sweep gas flow rate. Lower permeate pressure led tohigher water flux. Increasing the sweep gas flow rate increased waterflux, especially at low sweep flow rates (or low values of sweep/feedflow rate ratio). For example, as sweep/feed flow rate ratioincreased from 0% to 5.9% at a permeate pressure of 4.8 bar, waterflux increased by 2.5 times from 0.13 slpm/m2 to 0.47 slpm/m2.

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114112

The increase in water flux leveled off at a sweep/feed flow ratio of0.05–0.10; using a sweep gas flowing at only 5–10% of the feedflow rate provides the required water vapor removal for efficientmembrane dehydration.

Eq. (7) correlating partial pressure measurements and permeanceassumes perfect countercurrent flow in the module. However, themodule configuration [as shown in Fig. 5(b)] indicates the existenceof a non-ideal countercurrent flow pattern. The ChemCAD processsimulator, modified to include a sweep/countercurrent membraneunit operation, was used to understand the efficiency of the sweepmodule. Fig. 8(b) shows the ideality of membrane separation perfor-mance in the sweep module, which is defined as the percentage ofmeasured water flux relative to the simulated (ideal) water flux. Thevalues of simulated (or ideal) water flux were obtained assuming awater vapor permeance of 930 gpu. Mixed-gas permeance data usedfor other components were shown in Table 2. The ideality for waterflux was about 80% for the sweep module, and was essentiallyindependent of sweep flow rates. Considering the complexity of theflow distribution in the permeate channel and associated non-perfectcountercurrent flow, an ideality of 80% seems quite good. This

0.0

0.2

0.4

0.6

0.8

0 2 4 6 8 10Permeate pressure (bar)

Sweep/feedflow rate ratio = ~ 0.10

No sweep

Wat

er v

apor

flux

(slp

m/m

2 )

Fig. 7. Comparison of water vapor flux with and without sweep, as a function of

permeate pressure for a sweep/countercurrent module. Feed pressure was about

35 bar. The lines are provided to guide the eye.

0.0

0.2

0.4

0.6

0.8

0 0.1 0.2

Wat

er v

apor

flux

(slp

m/m

2 )

Sweep / feed flow rate ratio

Permeatepressure = 4.8 bar

7.9 bar

Fig. 8. (a) Effect of sweep/feed flow rate ratio on water vapor flux at different perme

contained 1100 ppm water at a flow rate of 0.050 m3(STP)/s; the sweep gas was dry

membrane module separation performance, which is defined as the percentage of meas

flux were obtained using a ChemCAD process simulator with assumed membrane wat

behavior was consistent with other studies showing that the sweepflow distribution pattern does not strongly affect the performance ofmodules that use a sweep gas [37].

The effect of sweep flow rates on water removal and methane lossare shown in Fig. 9. Fig. 9(a) shows that increasing sweep flow rateincreased water removal efficiency, while Fig. 9(b) shows thatmethane loss was essentially independent of the sweep flow ratestested. In other words, water flux in this module increased 5–10times by using a sweep gas, while the loss of hydrocarbons remainedalmost unchanged. This result is consistent with the membranestamp data shown in the Part I study [6], and confirms that processdesigns with sweep/countercurrent operation make membrane sys-tems more attractive for natural gas dehydration.

A water vapor permeance of 930 gpu can be obtained with currentPebaxs-based composite membranes, as demonstrated in the fieldtest [see Fig. 8(b)]. Membranes with higher water/methane selectivityvalues can be obtained by increasing the thickness of the selectivelayer, which would decrease CH4 permeance while having minimaleffect on water vapor permeance [6]. In future work, the membranesupport and gutter layers will be further optimized to reduce theirresistance to water transport.

The polyether phase in Pebaxs polymers is rubbery, and showshigher permeance to C2–C6 hydrocarbons than to methane (as shownin Table 2), thus avoiding the potential heavy hydrocarbon condensa-tion on the upstream side of the membrane. Condensation of heavyhydrocarbons can damage the membrane, and is usually avoided byheating the feed gas or removing heavy hydrocarbons from the gasbefore it is fed to the membrane modules, especially when mem-branes based on glassy polymer selective layers are used [38].Therefore, a membrane system containing Pebaxs-based membranescan be simple, with minimal pretreatment required. In addition, thePebaxs-based membranes only show moderate selectivity of higherhydrocarbons/methane. For example, the n-pentane/methane selec-tivity is only 7.4, as shown in Table 2. Therefore, high values forwater/heavy hydrocarbons selectivity can be expected, which wouldprovide low loss of heavy hydrocarbons in the membrane systemduring natural gas dehydration.

5. Conclusions

The successful development of membrane technology for naturalgas dehydration as a competitively viable alternative to conven-tional glycol dehydrators faces two substantial challenges. First,

0

20

40

60

80

100

0 0.1 0.2

Mea

sure

d/pr

edic

ted

wat

er fl

ux (%

)

Sweep / feed flow rate ratio

4.8 bar

Permeatepressure= 7.9 bar

ate pressures (4.8 bar and 7.9 bar) for a sweep/countercurrent module. The feed

(o10 ppm water) and contained mainly nitrogen and methane. (b) Ideality of

ured water flux relative to the simulated water flux. The values of simulated water

er permeance of 930 gpu. The lines are provided to guide the eye.

0

15

30

45

60

0 0.1 0.2

Wat

er re

mov

al (%

)

Sweep / feed flow rate ratio

Permeatepressure = 4.8 bar

7.9 bar

0

2

4

6

8

10

0 0.1 0.2

Met

hane

loss

(%)

Sweep / feed flow rate ratio

Permeatepressure = 4.8 bar

7.9 bar

Fig. 9. Effect of sweep/feed flow rate in a sweep module on (a) water removal efficiency and (b) methane loss. Water removal efficiency is defined as the percentage of feed

water permeating through the membrane. Methane loss is defined as the percentage of methane in the feed entering the permeate. Feed gas contained about 1000 ppm

water at a flow rate of about 0.050 m3(STP)/s. The sweep gas was a dry gas (o10 ppm water) containing mainly methane and nitrogen. The lines are provided to guide

the eye.

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114 113

when the polymers that are highly selective for water/methaneare made into industrial thin film composite membranes, themajority of the transport resistance to water can come from thesupporting microporous and paper layers, instead of the selectivelayer; this factor causes a significant decrease in the apparent waterpermeance and H2O/CH4 selectivity of the composite membranes.Second, the separation is restricted by moderate feed-to-permeatepressure ratios. The first challenge was addressed in Part I of thisstudy. The work described in this paper (Part II of the study)addressed the second challenge by investigating suitable processdesigns to overcome the limitations caused by the feed-to-permeatepressure ratios.

The membrane process simulation analysis showed that, amongfour process designs studied, a countercurrent design using a dry gasstream as sweep is the most economic process for natural gasdehydration. For example, a natural gas stream of 1.5 m3(STP)/s canbe dehydrated from 1000 ppm to 100 ppm with a membrane area assmall as 120 m2 and a potential methane loss of as little as 0.78%,which is very attractive compared to conventional glycol dehydrators(with a methane loss of 1%). The membrane has water vaporpermeance of 1000 gpu and H2O/CH4 selectivity of 500.

To verify the feasibility of the sweep/countercurrent design, a fieldtest of a sweep/countercurrent spiral-wound module was performedat a natural gas processing plant. The water vapor flux increased withdecreasing permeate pressure and increasing sweep flow rate. Themeasured water flux reached about 80% of the value calculated with aprocess simulation package for a perfect sweep/countercurrent mod-ule, confirming the good sweep efficiency of the sweep spiral-woundmodules.

In summary, this study shows that membrane processes with drygas sweep on the permeate can be competitive with conventionalglycol dehydrators for natural gas dehydration, and that spiral-woundmodules with good sweep/countercurrent efficiency can be made. Inthe future, the membrane configuration will be further optimized tominimize the water transport resistance in the microporous supportand gutter layers.

Acknowledgments

We gratefully acknowledge the financial support of this workby the United States Department of Energy Small Business Innova-tion Research (SBIR) program (Grant number DE-FG02-03ER83711).

We also thank Towne Exploration (Isleton, CA) for hosting thefield test.

Nomenclature

Am membrane area (cm2)gpu gas permeation unit [10�6 cm3(STP)/(cm2 s cm

Hg)]‘ thickness of the membrane selective layer (cm)NA steady state flow rate of gas A across the mem-

brane [cm3(STP)/s]NP total steady state flow rate across the membrane

[cm3(STP)/s]p1 membrane permeate pressure (bar)p2 membrane feed pressure (bar)(P/‘)A permeance of composite membrane to gas A (gpu)DpA,0 partial pressure difference across the membrane

for gas A at y¼0 (bar)DpA,L partial pressure difference across the membrane

for gas A at y¼L (bar)slpm standard liter per minuteslpm/m2 standard liter per minute per m2 membrane areax1 molar fraction in the permeate streamx2 molar fraction in the feed streamxF,A molar fraction of component A in the feed inlet gasxR,A molar fraction of component A in the residue gas

exiting the module

Greek letter

aA=B permeance selectivity of gas A over gas B

Subscript

1 membrane permeate side2 membrane feed sideA component AB component BF feed gas entering the membraneR residue gas exiting the membrane

H. Lin et al. / Journal of Membrane Science 432 (2013) 106–114114

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