RPSEA 

Field Assessment of Anti-Fouling Polymeric 

Membrane Coatings for Treatment of 

Barnett Shale Flow-Back Produced Water

Report No. 08122-05.05

Barnett and Appalachian Shale 

Water Management and Reuse Technologies

Contract 08122-05 

March 31, 2011

Principal Investigators Project Manager

Xiaofei Huang1, Dileep Agnihotri1, 

Benny D. Freeman2, Robert Hayes5, 

Sirirat Kasemset2, Albert Lee2, 

Hua Li1, Mukul Sharma3, and 

Steve Shiner4

Tom Hayes

Coordinator, Environmental 

Engineering Solutions

Gas Technology Institute

1700 S. Mount Prospect Rd. 

Des Plaines, IL

1 – Advanced Hydro Inc.

2 – Department of Chemical Engineering, The University of Texas at Austin

3 – Department of Petroleum and Geosystems Engineering, The University of Texas at Austin 

4 – Complete Filtration Resources, Inc.

5 – GeoPure Hydro Technologies
2

LEGAL NOTICE

This report was prepared by the The University of Texas at Austin as an account of work 

sponsored by the Research Partnership to Secure Energy for America, RPSEA. Neither RPSEA 

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3

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Summary

This report summarizes results of a brief (30 day) test of Anti-fouling Polymeric Membrane Coatings 

developed at the University of Texas at Austin (licensed and marketed by Advanced Hydro, Inc.) for 

treatment of shale flow-back “frac” produced water. Flow-back water collected from various shale wells 

in the Barnett Shale area was processed with ultra-filtration (UF) and reverse osmosis (RO) membranes. 

Membranes coated with polydopamine and PEG-amine were tested for their resistance to fouling due to 

organic and inorganic foulants in the feed water. Non-coated membranes were used as control for 

performance comparison.

Polyvinylidene Fluoride (PVDF) based spiral wound flat-sheet and polyacrylo nitrile (PAN) based 

outside-in hollow-fiber UF membranes were used in this pilot for pre-filtration for feed water prior to 

desalination using Polyamide RO membranes. Coated hollow-fiber UF elements outperformed noncoated elements, in terms of better cleaning efficacy (recovery) and higher flux, by almost 50%. In 

addition, coated spiral wound UF elements were easily maintained with daily chemical enhanced clean 

and fully recovered with regular acid-base cleaning operation. The coated spiral elements achieved 106 

hours run time without using a strong oxidant. The autopsy study revealed that UF membranes were 

fouled with organic substances, which could be mainly removed by acid/base clean. No significant flux 

loss was observed for both coated and non-coated seawater RO membranes during the pilot. Coated RO 

membrane demonstrated much better salt rejection than non-coated under high salinity conditions. This 

brief project provided validation of the coatings technology and its benefit for produced water 

purification. In addition, with improved salt rejection during desalination, a two stage nanofiltration or 

nano/RO membrane system with an optimized ultra-filtration (UF) or micro-filtration (MF) membrane is 

expected to enable membrane based purification of produced water while maximizing water productivity 

and minimizing maintenance and energy costs. 
5
6

Contents

1. Introduction 4

1.1 Background and Challenges of Shale Water Treatment and Reuse 4

1.2 Water Purification Technologies 5

1.3 Innovative Anti-fouling Polydopamine Membrane Coating Technology 7

1.4 Pilot Study Objective and Overview 8

2. Methods and Approaches 9

2.1 Pilot Site Location 9

2.2 Filtration Facility and System Design 10

2.3 Membrane Cleaning Procedures 12

2.4 Water Quality Analysis 12

2.5 Evaluation of Membrane Filtration Effectiveness 13

2.6 Investigation of Membrane Fouling 13

3. Results and Discussion 13

3.1 Water Quality Analysis 13

3.2 Ultra-Flo PAN Hollow Fiber UF Membrane Coated and Non-coated Results 16

3.3 Koch Coated PVDF Spiral UF Membrane Test Results 19

3.4 Hydranautics SWC3+ Coated and Non-coated Test Results 23

3.5 Membrane Autopsy Analysis 26

4. Conclusions and Recommendations 34

4.1 Pilot Summary 34

4.2 Key Notes form the Pilot 34

4.3 Future Work Recommendation 35

5. Acknowledgements 36

6. References 36
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1. Introduction

1.1 Background and Challenges of Shale Water Treatment and Reuse

The Barnett and Appalachian shale are among the largest and most active natural gas plays in the 

U.S. that geographically covers all or part of 20 counties in North Texas and large areas of Pennsylvania,

West Virginia, Ohio and New York. The Barnett and the Marcellus are among a number of natural gas 

shale basins in the U.S. The Barnett area is proven to have approximately 2.5 trillion cubic feet (59 km3) 

of natural gas reserves and is widely estimated to contain up to 27 trillion cubic feet (700 km3) of 

technically recoverable natural gas [1,2]. The Marcellus play in the Appalachian basin is considered to be 

larger in size and capacity in comparison to the Barnett. Both shale plays are considered to be 

unconventional gas formations and require hydraulic fracturing to make them productive. Hydraulic 

fracturing is a process in which large volumes of water, sand and chemicals are injected at high pressures 

into underground rock formations. The process creates fractures in the shale rock, allowing natural gas or 

oil to escape into the wellbore and be recovered.

Hydraulic fracturing (fracing), a necessary step for economic well performance, requires between 

1 and 4 million gallons (3,800 and 15,200 m3) of water for successful well completion. Vertical wells 

require approximately 1 million gallons (3,800 m3) and horizontal wells require 3-4 million gallons 

(11,400 to 15,200 m3) according to a recent survey among Barnett Shale Producers [3]. The same per well 

water use is found in Appalachian shales. Of the total water used by the industry, frac jobs represent about 

89 % and drilling represents 10 % [3]. This level of water demand poses a number of challenges for 

industry in the course of developing the Barnett and Appalachian shales for natural gas production. The 

industry must procure large volumes of water from regions where water may not be readily available. The 

typical shale gas well is a horizontal well that requires over 3.5 million gallons of water for completion 

(>80,000 bbls) requiring over 750 trucks to transport the water to the well site. This water is stored in a 

lined impoundment and stored there until the day the well is completed; during the well completion 

process, water from the impoundment is pumped down the well to achieve hydraulic fracturing of the 

shale rock. Following completion, about 1 million gallons of water is collected; this water contains 

elevated concentrations of salts, oils & greases, polyacrylamid polymers (added to reduce friction of the 

water along the well bore and in the cracks in the shale) and other soluble organics (such as volatile 

acids). Because well completion requires significant volumes of water within relatively narrow period 

of time (within a few days), water conservation has become an important element of water management

for the shale gas industry. 

Economics and stakeholder demands of water management are driving the oil and gas industry to

consider water reuse as a means to reduce the cost of transportation and to reduce the industry’s demand 

for fresh water supplies. One approach to achieving this goal is through the reuse of flowback and 

produced water; these water streams contain extreme levels of hardness, high scale formation potential, 

substantial presence of oils and greases, bacteria and polymers and other chemicals that can cause fouling 

on equipment involved in water storage, handling and treatment. If demineralization is required to meet 

the specifications for water reuse, challenges will include the control of factors that affect water 

processing. 
8

1.2 Water Demineralization Technologies

An important approach for decreasing shale gas industry water demand is to direct the flowback 

and produced waters collected from each completed well toward reuse in performing future well 

completions. The “reuse” option has the advantage of reducing the amount of freshwater required for 

each completion; other environmental benefits include reduced trucking requirements, easing of traffic 

congestion, reduced road wear, and decreased emissions and carbon footprint over the development life of 

the play. Generally, in the Barnett and Appalachian Shale Regions, the amounts of flowback water 

collected is usually in the range of 10 to 30 percent of the original water volume required for hydraulic 

fracturing. Some energy developers have treated flowback and produced water for the partial removal of 

suspended solids, oils and greases and scale forming constituents (e.g. calcium and barium) and blended 

these streams with freshwater for hydraulic fracturing at a ratio of more than 4:1; this practice can result 

in a blended water with a salt concentration of more than 10,000 mg/l which may require added expense 

for storage system integrity. Among some shale gas developers, concern still lingers over the potential 

effects of high dissolved solids in water on increased demands for friction reducers (though a new 

generation of friction reducers exhibits a higher tolerance for salt concentrations) and some operators are 

concerned over the long-term effects of high-TDS-water-mediated completions on shale gas well 

performance. For these reasons, many energy developers have been interested in the evaluation of 

technologies capable of an efficient recovery of demineralized water from brines for purposes of reuse. 

The main families of technologies capable of demineralization of brines include thermal 

processing and membrane treatment. One of the most reliable of separation technologies is the thermal 

distillation/condensation process which is a well proven water separations technology used in seawater 

desalination and in water management of many industries such as food processing (sugar), beverage, 

automotive, textile and commercial chemicals production. In the evaluation of this technology for water 

recovery from produced waters, the most commonly used evaporative treatment design has been the 

mechanical vapor recompression (MVR) system. This particular design is capable of handling a wide 

range of produced water salt concentrations from freshwater quality to above 200,000 mg/l, while 

achieving water recoveries of greater than 70%; recovered water is of distilled water quality, well suited 

for blending with freshwater supplies and suitable for storage in impoundments normally designed for 

holding freshwater. Design improvements have enabled modern MVR processes to operate at far higher 

energy efficiencies due to equipment features that efficiently recover heat from the condenser for transfer 

to the influent brines entering the evaporation chamber. This has enabled the process to operate with 

energy inputs as low as 0.11 mmBtu/barrel. The compelling advantage of MVR is its reliable 

performance under highly varied field conditions as evidenced by its ability to process brine streams of 

highly varied strength and quality over extended periods of time while generating a stream of distillate 

water for reuse and a concentrated brine containing nearly all of the salts that reduced in volume by more 

than 65%. 

Efficient demineralization of can also be achieved using membrane-based processes. Some 

energy companies have explored the use of reverse osmosis (RO) as a means of achieving brine volume 

reduction for shale gas waters and for conventional gas well produced water. For both types of produced 

water, pilot tests have shown that although a 3:1 reduction of brine volume could be achieved, and 

although a deionized product stream of good quality water (approaching distilled water quality) could be 

produced, there has been a practical limitation of RO to demineralize brines that exceed 40,000 mg/l TDS 
9

under the usual operating pressure drops of the RO process. Another challenge to membrane processes 

tested on produced waters in the past has been membrane fouling which has been among the typical 

operational problems noted in previous attempts to deploy the technology under actual field conditions. 

Fouling problems arise from the complex composition of the produced water and the effects of certain 

constituents on the membrane material. Free and dissolved oils collect on RO membranes causing them 

to lose their permeability and the presence of soluble hydrocarbons including volatile acids and BTEX 

can promote the growth of microbial films on ultrafiltration (UF) and RO membrane surfaces causing 

them to lose separation performance. These fouling problems are largely responsible for the lack of 

deployment success of RO in the oil and gas industry. Recent advances in applying rigorous 

preprocessing of produced water to remove fine suspended solids, oils & greases, and scale forming 

constituents show potential of reducing fouling problems, but extending membrane life is still a complex 

and site-specific problem. 

The development of a new generation of UF and RO membranes that could resist fouling while 

achieving energy savings in the demineralization would open even more opportunities to use low-energyinput membrane processes to recover a demineralized water stream from large volumes of moderately 

concentrated brines. In addition to providing a another tool for the recovery of product water from 

flowback waters that are of moderate strength (<50,000 mg/l TDS), membrane processing could provide a 

method of preconcentrating brines before feeding them into a thermal processor (such as an MVR) for 

further brine volume reduction and distilled water recovery. Whether used independently or together in a 

hybrid treatment system (as shown in Figure 1), the membrane treatment system and thermal processing 

represent important complimentary tools in the arsenal of water management for the shale gas industry. 

With regard to the improvement of membrane technology, recent research conducted under 

RPSEA support at the University of Texas has resulted in the identification and testing of polymeric 

coatings that, when applied to UF and RO membrane surfaces, enhance resistance to fouling and improve 

performance in salt separations in the treatment of synthetic mixtures designed to simulate shale gas 

flowback waters. Specifically, pPolydopamine (PDOPA) was found to be an effective anti-fouling 

surface coating for UF, NF, and RO membranes. The surface deposition of PDOPA improved the 

permeate flux for all types of membranes studied during oil/water emulsion filtration. For UF membranes, 

additional grafting by using poly(ethylene glycol) (PEG) further enhanced the fouling resistance. 

Significantly, the coated membranes exhibited substantially improved performance in terms of flux rates 

achieved when pressure drop across the UF and RO processes were kept constant. Potential energy 

savings of greater than 27% were estimated on the basis of these higher observed membrane flux rates as 

measured in the laboratory. 

On the basis of these results, a decision was made to initiate a field experiment to evaluate the 

effectiveness of the coatings technology to improve fouling resistance and performance in UF and RO 

treatment units under actual field conditions. To conduct this work, GTI has assembled a team including 

GeoPure (the owner of the UF/RO test skid), the University of Texas (owner of the coatings IP), and 

Advanced Hydro (membrane coatings deployment) to conduct a field evaluation of novel coated 

membranes with the cooperation of Devon Energy (the host for the demonstration). The purpose of this 

report is to document the field trials and results of this cooperative effort. 
10

1.3 Innovative Anti-fouling Polydopamine Membrane Coating Technology

Polydopamine is a polymer of dopamine (DOPA), which is a naturally-occurring hormone and 

neurotransmitter. Messersmith et al. [10] reported that dopamine solution will undergo polymerization 

under slightly alkaline conditions to form a polymer that mimics the properties of mussel adhesive 

proteins (specifically, M. edulis foot protein 5 (Mefp-5)). The mechanism of polymerization has been 

studied and elucidated elsewhere [11]. Messersmith’s suggested structure of polymerized dopamine or 

polydopamine (PDOPA) is shown in Figure 1. Polydopamine can non-selectively deposit onto virtually 

any surface. The deposition is on the order of 1-100 nanometers in thickness depending on contact time. 

The polydopamine layer exhibits excellent adhesion properties as it cannot be removed under sonication 

at extreme acidic conditions.

Figure 1: Dopamine and polydopamine structure. The exact structure of polydopamine is currently 

unknown; this structure was proposed by Messersmith et al. [10]

Polymerized dopamine coating for membranes has been developed for UF/MF, RO, and forward 

osmosis membranes by Freeman’s research group [12, 13]. Now this patent pending membrane coating 

technology is licensed through Advanced Hydro Inc. for commercialization. A PDOPA-coated membrane 

shows two advantages over a non-coated membrane: (1) The polydopamine layer is uncharged and 

hydrophilic, which should reduce oil/grease and protein/bacterial adhesion. (2)Polydopamine, under 

alkaline conditions, rearranges to form a quinone-like structure, providing a mean of further conjugation 

of other organic compounds.

The free hydroxyl groups on polydopamine can be used to covalently bind other molecules to the 

PDOPA-coated surface. McCloskey et al. [13] reported that amine-terminated poly(ethylene glycol) 

(PEG-NH2) can be grafted to a PDOPA-modified surface to improve the surface hydrophilicity without 

drastically altering the surface charge. Many other molecules could be grafted to the surface by a similar 

technique. Coating membranes with a thin layer of polydopamine should improve their resistance to 

fouling and provide a versatile starting point for further surface modification to a wide variety of 

membranes due to its nonspecific properties. Because the deposited polydopamine layer is so thin, the 

effect of the coating on membrane flux should be minimal.

PDOPA-PEG coating has been performed previously on finished UF and MF elements. Field 

pilot test results demonstrated superior permeation performance and fouling resisting property when used 

as seawater desalination pretreatment and waste water treatment. As shown by Figure 2, the attached PEG 

chain forms hair-structured layer, which is believed to be able to lock a water layer on membrane surface 

polydopamine

(proposed structure)
11

due to its super hydrophilic nature. Consequently, foulants cannot easily get close to membrane surface to 

form resilient fouling layer. Fouling layer would be more readily ripped off during cleaning operation. 

Figure 2: Cross-sectional structure of polydopamine and PEG coated membrane; conjugation of PEGNH2 to polydopamine surface at pH 8.8.

Strong oxidants (NaOCl, H2O2, ClO2, NH2Cl) are common cleaning reagents recommended by 

membrane manufacturers. With the PDOPA-PEG coating protection, membrane water productivity can 

be easily recovered by regular acid/base clean, even daily hot water clean can sustain the filtration 

performance for long period of time. This largely minimizes membrane operation and maintenance costs. 

Through coating treatment, membrane life can be extended by 30 percent, and 50 percent of cleaning 

chemicals can be saved. For the coated microfiltration and ultrafiltration membranes, energy saving is 

estimated to be 40 percent, and 10 percent reduction on labor cost as well as a 50 percent reduction in the 

cost of cleaning chemicals.

1.4 Pilot Study Objective and Overview

1.4.1 Objective

The overall objective of this pilot study was to evaluate the polydopamine treated membrane 

separation technology in the application of shale oil and gas produced water treatment and reuse. The 

scope of work included the following elements: (1) Evaluate the effectiveness of UF and RO membrane 

technology in treating oil and gas produced water; (2) Identify the improvements polydopamine coating 

technology can benefit the membrane procedure in terms of water productivity (flux-gallons per square 

foot per day) and fouling control; (3) Consolidate preliminary pilot experience, and propose an optimized 

future work plan.

1.4.2 Pilot Overview

The pilot was planned as a 30 day trial to recover shale flowback waters using ultrafiltration (UF)

and reverse osmosis (RO) processes. This short-term pilot was located in the Barnett Shale region near 

Fort Worth, TX, and co-existed with a vapor compression evaporation (VCE) treatment plant. Flowback 

water and produced water were transported from various production sites to the treatment facility and the 

treated demineralized water was trucked back to the fields to perform frac jobs. Raw water was treated 

by coagulation and sedimentation prior to processing by the membrane systems. Two types of 

ultrafiltration membranes were used in this pilot – the first set included two hollow fiber UF elements and 

the second set used two spiral wound UF elements. For the two hollow fiber UF elements, one was 

Polydopamine

Layer

PEG ad-layer

PDOPA-g-PEG polydopamine

(proposed structure)
12

treated with polydopamine and PEG, the other was not treated and used as “control” element for 

comparison purposes. Both elements were run in parallel simultaneously during the pilot. In addition, two 

spiral wound UF elements were treated with polydopamine and PEG and used in series inside a single 

pressure vessel. UF effluents from both systems were blended in a storage tank that was subsequently fed 

to the RO system. Two trains of RO membranes, each with three elements in series, desalinated the water 

processed by ultrafiltration. Elements in one of the RO train were treated with polydopamine. The two 

trains of the RO system were run alternately every 12 hours. 

Due to safety concerns, filtration systems were operated only during daylight hours. The pilot was 

also shut down during the Thanksgiving holidays for four days. During the normal running days, filtration 

was stopped for several times due to deficiency in raw water supply availability. As a result, the total 

accumulative filtration duration for spiral UF was 106 hours, hollow fiber UF was 58 hours, and about 50 

hours for each RO train. Pilot started from Nov. 17, 2010 and ended on Dec. 18, 2010.

2. Methods and Approaches

2.1 Pilot Site Location

The pilot system was located at the Maggie Spain Water Reclamation Facility operated under 

the support of Devon Energy in the Barnett Shale region near Fort Worth, TX. Raw produced water was 

trucked from different production locations and discharged into a lined storage pond before treatment. 

Raw produced water was primarily treated by coagulation and sedimentation before it flowed into 

compression evaporation facility. The pretreated water was pumped to a working buffer tank as supply to 

both filtration pilot systems. The reject from the UF and RO systems was pumped to the VCE treatment 

facility’s concentrate tank for off-site disposal. The permeate from the RO system was discharged to the 

VCE treatment facility’s condensate (finished water) lined pond. Figure 3 below demonstrates a layout of 

pilot facilities.

Figure 3: Layout of the pilot and collocated treatment facilities.
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2.2 Filtration Facility and System Design

Two sets of UF filtration systems were used in this pilot – the first one being a hollow-fiber, PLC

controlled ultrafiltration system with two hollow fiber UF elements in parallel. This system was designed 

and built by Advanced Hydro Inc. Due to high fouling potential of the produced feed water, hollow fiber 

modules were back-flushed every 15 minutes interval with air-scour for 30 seconds. To maintain the 

membrane performance, daily hot water clean (110 F) was also performed on hollow fiber modules at the 

beginning or end of each day’s operation. Among two hollow fiber modules, one was coated with 

polydopamine and PEG, while the other one was kept untreated as a control membrane. Filtered water 

from this system was stored in a common buffer tank before being fed to the RO system. Fully automated 

real-time flow and pressure data were recorded every minute by the PLC program. See Figure 4 for the 

layout of the filtration system.

The second UF system facilitated running of two spiral wound UF elements in series inside a 

stainless steel vessel. The second UF system also included two RO vessels; one vessel housed three 

dopamine-coated elements, with the other vessel containing three untreated RO elements. Both spiral 

wound UF elements were treated with polydopamine and PEG. This combined UF and RO system was 

designed and fabricated by Complete Filtration Resources and provided by GeoPure HydroTechnologies

and required manual operation and data logging. Membrane performance data and operation parameters 

were recorded manually every 20 minutes. Permeate from UF system was stored in a common buffer tank 

for feeding the RO system. UF vessel was flushed every 4 hours with heated caustic solution to maintain 

sustainable performance. A set of 5 micron cartridge filters were installed in front of UF segment to 

prevent large particles entering spiral wound membranes. These filters were removed after the first few 

days of the pilot study in order to allow the spiral and dead-end UF membranes to see the same 

“unfiltered” feedstock. A schematic drawing was illustrated in Figure 4 for the filtration system design.

Figure 4: Schematic of membrane filtration design

Feed

Feed

Reject

Ai

r

UF filtrated 

water 

storage

Reject

Cartridge filter

Permeate

Permeate

Permeate

Middle 

buffer tank

Reject

Reject

Coated RO

Coated UF

Non-coated

RO

Non-coated Coated UF

UF
14

Figure 5: Picture of the hollow fiber, spiral wound and RO filtration systems

After a few days of test runs, the optimal operating conditions were determined. Hollow fiber UF 

elements were operated dead end filtration mode with an initial feed flow rate of 5.8 GPM for each 

module. The backflush frequency was set to 15 minutes and its duration was 30 seconds. Spiral UF 

elements were operated in cross flow mode. The initial feed flow rate was set at 50 GPM for two 

elements, 45% of flow was recycled back to the feed tank to enhance the recovery, and 16% of flow was 

bled off as concentrate. The average feed pressure was about 55 PSI at the beginning, and initial permeate 

flow rate was set to about 20 GPM. When the permeation rate decreased to 10 GPM, caustic enhanced hot 

water cleaning was performed to recover flux. RO elements were run at feed flow rate of 15 GPM per 

vessel. The recovery (ratio of permeate to feed) of the RO system was kept low both to control operating 

pressures to less than 1,000 psi and to minimize scaling due to high feed hardness concentrations. The UF 

permeate buffer tank was used to maintain feedstock to the RO while cleaning was performed on the UF 

system.

Membrane Specifications

a. Hollow fiber UF membrane: Ultra-Flo U630C (Ultra-Flo Pte Ltd, Singapore) was chosen as the 

demonstration UF membrane. The original membrane fibers are made out of hydrophilic 

modified PAN (poly(acrylonitrile)). The fiber outer diameter is 1 mm, and the inner diameter is 

0.5 mm. The membrane surface area is 323 ft2. Recommended feed water turbidity is less than 5 

NTU. Typical product turbidity is below 0.1 NTU. It is an outside-in capillary style membrane. 

The pore sizes of this membrane ranged between 0.1 to 0.01 microns. This system was brought 

in with last minute changes because the Geopure skid did not support side-by-side testing for UF.

This limited our selection of membranes available at the time, with the selected membrane not 

optimum for use on produced water. However, we used the opportunity to generate data with 

side-by-side coated and non-coated control elements.

b. Spiral wound UF membrane: Sanitary ultrafiltration spiral membrane 8038 HFM-100 (Koch, 

Massachusetts, USA) was selected as another demonstration UF membrane. This membrane is 
15

fabricated from PVDF (poly(vinylidene difluoride)). The molecular weight cut-off of HFM-100 is 

50,000 daltons. The surface area of this membrane is 319 ft2. 

c. RO membrane: Seawater RO (SWRO) membrane SWC3+ (Hydranautics, California, USA) was 

selected as demonstration RO membrane. This membrane is a typical interfacial composite 

polyamide RO membrane with a nominal surface area of 400 ft2. Maximum feed water turbidity 

is required to be less than 1 NTU, and maximum cross flow must be less than 75 GPM. 

2.3 Membrane Cleaning Procedures

Two types of membrane cleaning methods were employed to recover UF membrane performance: 

chemical enhanced hot water clean (HWC) and standard chemical clean-in-place (CIP). 

HWC was performed daily to recover ultrafiltration membrane flux. For this procedure, RO 

filtered water was used as the cleaning solution. The membrane elements were first flushed with RO 

water to dilute the feed solution (produced/flow-back water) retained in the module. The water was heated 

to 125 oF. Caustic or citric acid was used to enhance the cleaning efficiency. Solution pH was adjusted to 

11.0 with caustic during alkaline cleaning for organic foulant removal, and adjusted to a pH of 2.0 with 

citric acid during acid cleaning for scale removal. This water was circulated through the membranes for 

30 minutes. After the HWC, membrane modules were fully flushed with RO water to remove the residue 

chemicals. Spiral membranes were flushed through the feed side, while backflushing was performed on 

hollow fiber membranes to remove foulants and residual chemical.

CIP was occasionally performed on ultrafiltration membranes as a more intensive cleaning

procedure. During CIP cleaning, caustic clean with NaOH (1% wt) was performed first for 2 hours and 

followed with citric acid (2%) cleaning for 1.5 hours. Cleaning solutions were also maintained at 125 oF 

during circulation. The membranes were fully flushed with RO water between each solution change. 

2.4 Water Quality Analysis

Water samples were collected and sent on different days to the UT-Austin laboratories for 

analysis where TOC and major component ions were analyzed. Water quality information related to feed 

water turbidity, RO feed, permeate conductivity, pH, temperature were collected in real-time at 20 minute 

intervals during membrane operation. In addition, alkalinity, hardness, calcium and chloride 

concentrations were tested on- site frequently using portable drop count titration test kits from Hach, 

Colorado, USA. 

Metal and non-metal elements (i.e. B, Na, Mg, Si, P, K, Ca, Fe, Sr, and Ba) in the water were also 

analyzed using inductively coupled plasma mass spectrometry (ICP-MS) (Agilent, model 7500ce) at 

Quadrupole ICP-MS Laboratory in Jackson School of Geosciences, University of Texas at Austin. Prior 

to testing by ICP-MS, the water samples were filtered with 5 um particle filters to remove any particulates 

in the water. Subsequently, each sample was diluted to 1:100 and 1:500 in 2% HNO3 to ensure

repeatability and to match with an operating range and a matrix of standard solutions.

Anions (i.e. Cl-, NO3

-

, SO4

2-

, and PO4

3-

) were analyzed by ion chromatography (IC). The water 

samples were diluted to 1:2000 in deionized water to ensure that the concentrations fell within the

detection limit of the equipment.
16

Total organic carbon (TOC) analyzer (Shimadzu, model TOC-VCSH) was used to analyze the 

organic carbon present in the water samples. In this analysis, non-purgeable organic carbon (NPOC)

method was used for analysis of the TOC content.

Total petroleum hydrocarbon (TPH) concentrations were analyzed using a silica gel treated nhexane extractable material (SGT-HEM) method of extraction and gravimetry (EPA 1664A). The TPH 

test was performed by TestAmerica Lab, an environmental testing firm in Pittsburgh. 

2.5 Evaluation of Membrane Filtration Effectiveness

Ultrafiltration membrane efficacy was evaluated through membrane permeation flux, applied 

pressure, and cleaning frequency. These data and parameters were recorded during the filtration 

operation. Coated membrane performances were tracked and compared to their non-coated control twins 

and results are summarized in section 3. 

Reverse osmosis membrane performance was assessed based on salt rejection and permeation 

flux. Since it is a pressure-driven procedure, besides the permeation flux data, the overall productivity 

was also evaluated in terms of recovery and net driving pressure. Salt rejection and temperature corrected 

permeate flux were calculated using an Excel based “RO data normalization utility” provided by 

membrane manufacturer Hydranautics. Calculation equations and conversion methods are discussed in 

the technical manual from manufacturer’s website [14]. 

2.6 Investigation of Membrane Fouling

Used elements were autopsied to investigate fouling following the 30 days of pilot work. Leading 

spiral UF membrane, tail RO membranes and both hollow fiber UF elements were dissected for further 

examination. Multiple membrane samples were extracted and examined for foulants. Control samples 

were prepared by acid and base cleaning and served as reference for clean membrane surface. ATR-FTIR 

and SEM measurements were performed at the UT-Austin to study the presence of foulants. 

3. Results and Discussion

3.1 Water Quality Analysis

Raw water quality data is summarized in Table 1. Turbidity, TDS, temperature, and pH 

parameters were tracked and recorded during the entire pilot and are presented at charts below. Few water 

samples for feed and permeates were collected for TOC and TPH tests. Alkalinity and hardness were 

analyzed.

Table 1: Summary of feed water quality
17

A wide variation in feed water quality was observed during the short pilot duration. This included 

feed turbidity ranging from 1.5 to 53 NTU with an average of 8.8. TOC varied from 77 to 95 mg/L with 

an average of 89 mg/L. TDS had a significant variation that directly impacted RO productivity. The

overall TDS varied from 14,000 to 63,000 mg/L with an average of 40,000 mg/L. Since the water went 

through pretreatment, its pH remained mostly acidic (due to addition of HCl to reduce carbonate scaling) 

with an average pH of 4.8. Hardness measurements based on calcium carbonates also varied significantly 

with a range of 3,500 to 12,000 mg/L with an average of 9,500 mg/L. Since this pilot was run during 

November and December timeframe in Dallas area, temperature variations also showed trends with 

lowest in morning and generally increasing throughout the day. 

All turbidity data (Figure 6), total dissolved solids (Figure 7) and temperature (Figure 8) are recorded in 

the figures below.

Figure 6: Feed water turbidity during the 30 day pilot

Figure 7: TDS variations for feed water during the 30 day pilot
18

Figure 8: Temperature variations for feed water during the 30 day pilot

Major component ions of collected water samples were analyzed using IC and ICP-MS methods. 

Only one water sample was collected from each stage of the filtration process. Measured values of these 

ions are listed in Table 2. All the analyzed ions were retained in the UF permeates. There were gain and 

loss on the reduction percentage of UF permeate from the feed ionic concentrations, which was mainly 

due to the fluctuation of feed water quality. Rejection ratios for most tested ions were high above 99% on 

RO membrane. About 80% removal of Boron and 75% removal of nitrate achieved by this single stage 

SWRO membrane. Only uncoated RO permeate water was accepted full span ion analysis. The 

comparison data for ion rejections (mainly on hardness ions and chloride) for coated and non-coated RO 

membranes is discussed in section 3.4. No TPH was detected for the collected feed and permeate water 

samples. We suspect that water analysis for TPH was not accomplished within the required timeframe to 

get representative data.

Table 2: Ion analysis results for raw feed water, UF permeate, RO permeate, and RO concentrate.
19

3.2 PAN Hollow Fiber UF Membrane: Coated vs. Non-coated Test Results

Hollow fiber ultrafiltration membranes were tested in parallel for PDOPA/PEG coated and noncoated elements. All operating conditions and cleaning procedures were identical for both elements. 

During the 60 hour run, coated membranes demonstrated significantly better performance than noncoated membranes. Coated elements generated more permeate (higher flux) than non-coated elements. 

This advantage in water productivity became more significant in the latter stages of running as 

membranes fouled. Figure 9 shows the permeate water flux measured during the pilot. The membranes 

we used were available off-the-shelf and typically used for water treatment with a recommended 

maximum feed turbidity of less than 5 NTU. Such membranes are not optimum for produced water pretreatment; however, these were the only membranes available for the side-by-side testing. 

Pilot testing began with new membranes with the initial flux for both hollow fiber modules set at 

25 gallons per square foot of membrane per day (GFD). Flux for both membranes dropped gradually over 

time. Although both coated and non-coated control membranes lost flux during the 60 hours of pilot 

operation, coated membrane provided nearly 50% higher flux compared to that of a non-coated control. 

Irreversible losses from 25 to 15 GFD are due to the fact that membrane module is not the optimum for 

high turbidity water and had a very tight fiber geometry to achieve higher surface area. We found 

corroborating evidence during membrane autopsy and there were significant amounts of large particulate 

residues retained inside the module that did not get removed during the backflush processes.

Figure 9: Ultrafiltration Membrane permeation flux for 3500 minutes; HWC1: caustic enhanced hot water clean, 

pH=11, 30 minutes circulation at 110 oF. CIP: clean-in-place, with caustic clean first, pH=11, 2 hours circulation at 

HWC 1

HWC 1

HWC 1

RO water 

back flush

CIP
20

110 oF; followed with citric acid clean, pH=2, 1.5 hours circulation at 110 oF. RO Water back flush: UF 

membranes were flushed using treated demineralized water from the RO process. 

During the operation, we performed chemical cleaning on both elements as documented in Fig 9. 

Caustic enhanced hot water cleaning was found to be effective to recover flux, since organic fouling was 

the main issue for the feed water.

During the final days of pilot, the coated membrane was producing approximately 50% more 

water than the non-coated control. In addition, “hot water clean” (HWC) and “clean in place” (CIP)

cleaning recovered coated membrane flux more effectively than the non-coated membrane. The typical 

flux for the coated membrane was maintained between 12-16 GFD using only daily caustic enhanced 

HWC. On the other side, caustic HWC and acid/base CIP barely retrieved flux on non-coated control 

membrane, suggesting formation of resilient fouling layer for acid and base cleaning chemistries. To 

remove such foulant, harsh chemicals (perhaps strong oxidants such as bleach or hydrogen peroxide) may 

be necessary to recover membrane performance. The typical flux for the non-coated UF membrane was 

maintained between 7-11 GFD.

Trans-membrane pressure (TMP) is a measure of the pressure difference between the feed side 

and the permeate side of the membrane during the filtration process. To create permeation across the 

membrane, TMP acts as the driving force for the filtration processes. It also directly relates to pore 

blocking (fouling) and energy consumption for the separation process. When a fouling layer is present, 

the membrane becomes less permeable, leading to an increase in TMP values. We measured TMP for 

both coated and non-coated membranes during the pilot side-by-side testing as represented in Figure 10. 

The non-coated membrane consistently showed higher TMP (about 25%) compared to the coated 

membrane. The typical differences are of the order of 4-5 psi. Such data is consistent with the earlier 

piloting of this technology in seawater desalination and laboratory development at laboratory scale. The 

lower TMP results from hydrophilic PDOPA + PEG coatings and their propensity to prevent adhesion of 

the foulants to the surface. 

Figure 10: Trans-membrane pressure data for two side-by-side UF hollow-fiber elements. The red data 

represent the coated membrane and blue data represents non-coated control element.
21

Despite nearly 25% lower TMP across the coated filtration membrane, this membrane was 

generating about 40-50% higher flux during the second half of the pilot (post initial fouling). The 

combined effect of such data can be quantified as specific flux (GFD/psi). The ratios of flux and TMP 

(Specific flux) are plotted in Figure 11. This data represents a direct measurement of productivity and 

energy consumption for the filtration process. Notably, higher specific flux is obtained for the coated 

membrane. The coated membrane operated at about 1.5 GFD/psi specific flux compared to non-coated 

operating at 0.8 GFD/psi, nearly a factor of 2X higher. The coated module’s performance increased over 

time.

Figure 11: Specific flux data for two side-by-side UF hollow-fiber elements (top). The red data represent 

the coated membrane and blue data represents non-coated control element. A ratio of two specific fluxes 

comparing coated and non-coated is also plotted (bottom).
22

3.3 Koch Coated PVDF Spiral UF Membrane Test Results

Two spiral UF elements were coated with PDOPA/PEG and tested in this pilot for about 6500 

minutes. During the initial test runs, we established a requirement for HWC cycle every 4 hours to 

maintain a sustainable flux for subsequent feed to the RO stage. However, at later stages, although the 

water was more turbid, after initial fouling, the membranes reached stable operation requiring HWC once 

or twice a day during 12 hours of operations. During the 30 day period, we performed a total of 3 CIP 

cycles for this set of membranes and nearly 20 HWC processes to manage fouling. We expect to improve 

this further with an optimized pre-filtration cartridge that can capture oily components (see discussion 

later).

After treated with polydopamine/PEG coating, Koch spiral 8038 HFM-100 membranes can 

sustainably treat shale produced water, although frequent caustic cleanings were required. The average 

filtration productivity was 35 GFD for the 106 hours of runtime, which is relatively large for a spiral UF 

module under this feed water condition. Trans-membrane pressure increased gradually over the test 

duration. Apparently, the cleaning could not fully restore the original separation performance in terms of 

energy consumption. Although the flux mostly recovered after each cleaning cycle, feed pressure kept 

increasing over time, which results in the 50% loss on specific flux from its original value. This indicates 

the adsorption of foulants on the membrane surface. However, permeate flux could be effectively 

recovered on coated membrane by caustic HWC. The cleaning temperature was found to be critical to 

restore the permeate flux. Cleaning temperatures around 120oF provided an excellent recovery of the 

membrane flux during circulation of caustic solution, which is within the cleaning temperature range 

required by the membrane manufacturer (105-130oF) [15]. The entire data set is plotted in Figure 12 

below as flux (top), TMP (middle) and specific flux (bottom). For this set of membranes, we did not have 

a control element due to skid design limitations. Based on achieving a sustainable flux of over 35 GFD, 

we are clearly seeing a significant benefit due to the presence of the coating layer. For such challenging 

water, it may be difficult to achieve flux in the range of 25 GFD without the coating layer for this kind of 

membrane.
23

HWC 1 HWC 1

HWC 1

CIP CIP CIP
24

Figure 12: Spiral UF module pilot data. Top: membrane permeation flux for 6500 minutes; Middle: transmembrane pressure for 6500 minutes; Bottom: specific flux (Flux/TMP) for 6500 minutes. HWC1: 

caustic enhanced hot water clean, pH=11, 30 minutes circulation at 110 oF. CIP: clean-in-place, which 

uses caustic clean first, pH=11, 2 hours circulation at 125 oF; followed with citric acid clean, pH=2, 1.5 

hours circulation at 125 oF.

Since, there is no equivalent non-coated data we can be use to compare the performance, we 

studied some data from previous pilot tests to get basic ideas of how such membranes may perform 

without coating in treating produced water. However, it must be noted that the data we present below, 

corresponds to water that was very different and much better quality compared to the water in current 

pilot. Only basic raw water quality data is available from this old pilot. Table 3 below illustrates the 

difference between water quality for our pilot at the Maggie Spain Facility near Fort Worth (Fort Worth 

pilot) and the water handled by an old pilot system previously operated elsewhere (both pilot efforts 

utilized the same spiral UF membrane).

Table 3. Feed water qualities of Fort Worth pilot and a pilot conducted elsewhere using the same spiral 

UF membrane (Old pilot).
25

It is clear that in the Fort Worth pilot we had an average of 60% higher turbidity and nearly 2x 

higher TDS. The old pilot had better feed water quality than Fort Worth pilot. Five test runs were 

conducted during the old pilot, with each test run lasting for about 3 hours. Total running time was about 

1000 minutes. In Figure 13, the first 1000 minutes operation results from Fort Worth pilot are plotted and 

compared with the other pilot data. 

Figure 13: Pilot data of coated spiral UF membrane plotted with pilot data of non-coated spiral UF 

membrane. Left: Water flux for the first 1000 minutes run; Right: Specific flux (Flux/TMP) for the first 

1000 minutes run. 

As seen from Figure 13, non-coated Koch spiral UF membrane provided stable flux at around 40 

GFD. This performance was also maintained by caustic enhanced HWC. It was reported that caustic clean 

was conducted in between each test run. Although the Fort Worth pilot featured feed water with more 

challenging features (higher suspended solids and higher salinity), the coated spiral UF membrane 

displayed comparable performance as a non-coated membrane treating less challenging water. The flux 

change was more dramatic for the Fort Worth pilot, but the coated membrane demonstrated a recoverable 

surface in the presence of a highly fouling feed stream.

Although this data is not a direct comparison, it is able to shed light on the fact that with a water 

that has nearly 2x higher TDS and 80-90 mg/L organics and 60% higher turbidity, coated membranes can

achieve about 35 GFD flux over a nearly 6x longer pilot duration, exhibiting a very recoverable operation. 

We attribute such benefits to the use of the coating applied to the membranes. 

During the CIP cleaning of Koch spiral UF membranes, significant amounts of oil emulsion 

residue were found in the circulated citric acid based cleaning solution (Figure 14, left) during the 

beginning of the cleaning cycle. However, this residue disappeared during the latter half of the circulation 

process. Whenever this happened, the feed pressure could not be effectively recovered by cleaning, 

suggesting the oil residue was being entrapped against the inside of the membrane (mostly inlet) during 

the cleaning circulation mode. To resolve this problem, a cartridge filter needs to be installed in-line

before the cleaning solution goes back into the membrane vessel; this filter was not a component of the 

existing skid. 
26

The reject solution from the caustic cleaning was very dirty and dark in color, which suggests that 

most foulants were organic substances and effectively removed during the caustic cleaning. 

Figure 14: Oil residue/particulates collected during citric acid cleaning (left), and a dark solution collected 

during caustic cleaning (right) representing predominantly organic foulants adhering to the membrane.

3.4 Reverse Osmosis Membrane Performance 

Water processed (permeate) from the ultrafiltration membranes was collected in a common tank 

and supplied to the RO as feed for subsequent removal of dissolved ions (desalination). The feed flow rate 

was fixed at 15 GPM with an initial recovery at 50%. Antiscalant Flocon 260 was added into the feed line 

to prevent scaling in the concentrate side of the RO elements. As discussed earlier in Figure 7, TDS 

values fluctuated greatly from day to day. The highest measured value for TDS reached about 63,000 

mg/L. Since the maximum capacity of the RO feed pump was limited to 1000 PSI, during high TDS 

conditions, recovery could only be maintained at 10-15% due to the high salinity of the feed. Specific 

permeate flux and rejection were normalized by temperature and plotted in Figure 15, for which, all the 

flux and rejection data were temperature corrected to 25oC.

As seen below in Figure 15 (top), relatively, no significant specific permeate flux changes were 

noticed on both coated and non-coated SWC3+ membranes. The normalized specific flux data show 

significant random variations for later half of the pilot when feed TDS was more than 50,000 mg/L. This 

may appear unusual; however, such behavior is generally attributed to significant drop of net driving 

pressure across the RO membrane. ssociated with understood in terms of net driving. This reduced net 

driving pressure resulted from significantly increased TDS values. As seen in Figure 16, feed water TDS 

increased about 2.4-3 times in the latter half period of pilot run, which leads to the reduction on the 

membrane recovery as the net driving pressure dropped. The limited maximum feed pressure led to the 

loss of net driving pressure due to the increase on osmotic pressure of the higher saline feed. This

membrane material gets more permeable at lower net driving pressure, which might be due to the less 

compacted structure of membrane sheet at low pressure range. 
27

Figure 15: Coated and non-coated SWC3+ pilot test results. Top: Normalized flux; Bottom: Normalized rejection.

During our testing, we found higher and more stable salt rejection for coated membranes. The 

non-coated membranes experienced a significant decrease in rejection when the feed TDS increased. 

Rejection started for non-coated membranes at 99.5% and dropped to the lowest value of 96.6%. 

However, for coated membranes, rejection remained well maintained above 99.5% throughout the pilot. 

This is expected behavior for polydopamine coating on RO membranes. Based on this observation, we 

expect a system where nano-filtration membranes with polydopamine coating may perform very well for 

filtration of high TDS water streams and eventually enable either a NF-NF or NF-RO combination of 

desalination for a wide range of salinity streams. We recommend that such a study be performed in a fullscale pilot effort in the near future.
28

Figure 16: Top: Net driving pressure (NDP= P- ; Bottom: Feed water TDS values vs. membrane recovery 

during the RO pilot test.

Chloride and hardness (Ca2+, Mg2+) concentrations were measured using on-site titration for feed 

and permeate RO water to evaluate the individual rejection of specific ions. Results are presented in Table 

4. Rejection of chloride and hardness ions was higher for the coated membrane than for the non-coated 

membrane. This observation is also in agreement with the rejection data presented in Figure 15 (bottom) 

earlier, which evaluated from the on-line conductivity measurement. 
29

Table 4: Specific ion rejection by coated and non-coated SWC3+ membrane

3.5 Membrane Autopsy Analysis

Leading spiral UF elements, tail RO elements for coated and non-coated membranes, and two 

hollow fiber UF elements were opened for fouling evaluation. The resulting autopsy examination and 

analysis are discussed in the following sections.

3.5.1 Membrane Leaf Examination

During membrane autopsy, we made following observations (Figure 17 below) –

1. Coated Spiral UF: Membrane leaves were dark in color, and surface looks oily. Brown stains were 

found on brine spacer leaves.

2. Non-coated SWC3+: Membrane leaves were light pink in color, and surface looks clean.

3. Coated SWC3+: Membrane leaves were light brown in color, and surface looks clean.

4. Non-coated hollow fiber UF: Fibers were brown in color. 

5. Coated hollow fiber UF: Fibers were dark brown in color.

Spiral UF before autopsy RO Before autopsy
30

Spiral UF element leading edge RO element leading edge

Spiral UF membrane leaf Non-coated RO membrane leaf

Spiral UF brine spacer Coated RO membrane leaf

Figure 17: Opened elements leaf examination.
31

3.5.2 Hollow fiber UF Membrane Examination

Fibers were cut from the opened modules and studied. Sample fibers were cleaned with either 

acid/base (washed) or acid/based and bleach (washed and bleached) with sodium hypochlorite. Images of 

fibers as obtained from the elements without cleaning, post washed and post washed and bleach were 

taken and are shown in Figure 18. Coated fibers were brown in color as expected from the coating 

polymer. Fouled fibers in both coated and non-coated elements are slightly darker in color. The fiber 

color became lighter after cleaning (washed). Cleaning with a strong oxidant (bleach) restored the 

original color of the membrane fiber. After bleaching, coated fibers turned white in color, which suggests 

the removal of the polydopamine/PEG coating layer. In addition, coated fibers exposed native surface 

while uncoated fibers retained some color, most likely from unrecoverable fouling.

Figure 18: Coated and non-coated fibers treated with different cleaning reagents. Left: uncoated fiber, from bottom 

up: untreated, washed with acid /base, washed with acid/base and bleached; Right: coated fiber, from bottom up: 

untreated, washed with acid/base, washed with acid/base and bleached.

3.5.3 Attenuated Total Reflection Fourier Transform Infrared Spectroscopy 

(ATR-FTIR) Examination of Membranes

Membrane samples taken from the opened elements were cleaned and compared with unclean 

raw samples using Attenuated Total Reflection Fourier Transform Infrared Spectroscopy (ATR-FTIR). 

After cleaning, the virgin surface was mostly observed. Therefore, in Figure 19-22, cleaned membrane 

spectra should represent the pristine surface of either coated or uncoated membranes (probably mixed 

with a little foulants depending on how well the membrane was cleaned).
32

Figure 19 shows IR spectra of cleaned spiral UF membrane and fouled UF membrane. The 1500-

650 cm-1 region contains characteristic peaks of PVDF [16]. Peaks in this region are attenuated in the 

fouled membrane, indicating that the original surface properties were masked by the fouling layer. The 

broad band centered at 3300 cm-1is due to the overlapping of vibration of hydroxyl (-OH) groups from 

the coating layer and carboxylic groups from the fouling layer [17]. Small peaks in the range of 1550-

1650 cm-1of the fouled membrane are probably due to the contribution of amide (C=O or protonated 

amine), nitro (N=O or R-NO2) groups and carboxylic functional groups [18, 19]. The increased 

absorbance at around 1000 cm-1 may suggest the presence of polysaccharide-like substances. The peak 

between 900-1100 cm-1probably corresponds to a ketone group of polysaccharide in cellulose [20]. This 

peak overlapped with attenuated PVDF peaks close by and formed a region of increased absorption at 

around 1000 cm-1in the fouled membrane spectra. The differences between the spectra of the fouled 

membrane and the cleaned membrane suggest that organic matter is the primary foulant on these spiral 

UF membranes. 

Figure 19: IR spectra for coated spiral UF membrane, cleaned versus fouled.

Figure 20 below presents IR spectra for cleaned coated and non-coated SWC3+ membranes. The 

peaks at 1540, 1610, and 1660 cm-1correspond to C-N stretching (amide II), aromatic ring breathing, and 

C=O (amide I) functional groups which are associated with the presence of amide groups [21]. 

Membrane IR profiles are also characterized by a broad band centered at 3300 cm-1, which corresponds to 

the presence of amide (N-H), carboxylic (COOH), and hydroxyl (-OH) functional groups. Theses 

polyamide characteristic peaks were all slightly attenuated for the coated SWC3+ spectrum, which 

indicates the presence of the coating layer on the top of the virgin surface.
33

Figure 20: IR spectra for non-coated and coated RO membranes.

When comparing the fouled membrane with the clean membrane, as shown in Figure 21 for a 

non-coated membrane and in Figure 22 for a coated membrane, characteristic polyamide peaks in the 

spectrum were attenuated by various amounts compared to the spectrum of a clean membrane. This result 

implies a thin layer of bounded foulants or coating, which hindered the penetration of infrared radiation 

into the virgin polyamide layer. The extent of peak reduction semi-quantitatively suggests the thickness of 

fouling layer or coating layer.

Figure 21: IR spectra for non-coated RO membrane, cleaned versus fouled.
34

As seen in Figure 21, uncoated membranes show attenuation due to the presence of the foulant 

layer. Although no noticeable performance decrease was observed with these membranes, the infrared 

results suggest early signatures of fouling; 30 days of service were apparently not enough to produce a 

strong impact of fouling on flux.

Figure 22: IR spectra for coated RO membrane, cleaned versus fouled.

As seen in Figure 22 above, there is very little change in FTIR intensities on coated membrane 

compared to that of the coated and cleaned membrane, suggesting evidence of a much thinner layer of 

adsorbed foulants. These intensity changes are significantly less than those of a non-coated membrane,

suggesting that the coating has significant anti-fouling impact.

3.5.4 SEM Examination

SEM images were taken for RO and spiral UF membranes obtained during the autopsy of the 

membrane elements. SEM pictures of smaller to larger magnification are arranged top down in Figure 23

and 24. Both coated and non-coated SWC3+ membranes retained their original surface morphology. The 

surface looks clean for these RO membranes. No specific deposits were found on the scanned samples. 

On the other hand, the coated spiral UF membrane (8038 HFM-100) surface was fully covered with 

colloidal clusters, which were composed of small nodular lumps about 100 nm in size. The original 

porous structure of the PVDF membrane was totally covered by the fouling layer. This fouling layer was 

mostly removed by CIP cleaning using caustic and citric cleaning mediums. The FTIR data very much 

corroborates the SEM examination.
35

 Non-coated SWC3+ Coated SWC3+ 

Figure 23: SEM images for non-coated RO (left) and coated RO (right) at various magnification 

levels (1.3k, 4.5k, 40k, top down)
36

Coated 8038 HFM-100 Coated-Cleaned 8038 HFM-100

Figure 24: SEM images for coated 8038-HFM UF membrane. Left column represents fouled and non-clean 

membrane surface. Right column represent fouled membrane after a clean process that included several hours of 

wash with caustic and citric. Top down represent various magnification levels (1.3k, 4.5k, 40k).
37

4. Conclusions and Recommendations

4.1 Pilot Summary

PDOPA coating technology was tested on hollow fiber ultrafiltration PAN membranes (U630C, 

Ultra-Flo), spiral ultrafiltration PVDF membranes (8038 HFM-100, Koch), and seawater reverse osmosis 

membranes (SWC3+, Hydranautics) in the application of treating shale “frac” flowback water. Parallel 

tests were performed on hollow fiber UF membranes with coated and non-coated elements. Parallel tests 

for SWRO membranes were conducted by running coated and non-coated elements alternately. Only 

coated spiral UF membranes were tested in this pilot due to skid limitations.

Feed water supplied to the membrane systems was a blend of shale gas frac flow-back and 

produced water pretreated via coagulation/flocculation and gravitational settling. This water was rich in 

organic substances (soluble and insoluble), suspended solids, hardness, barium, sulfate, salinity, bacteria 

and was acidic in nature. Raw water quality fluctuated greatly during the entire pilot time span.

For the parallel tests on hollow fiber UF and SWRO, coated membranes demonstrated improved 

performances over their virgin non-coated counterparts without any loss of flux due to coating. Coated 

hollow fiber UF membranes gave about 50% more water productivity than the non-coated ones. Cleaning 

operations were more efficient in recovering flux and pressure for the coated membranes. Coated 

seawater RO membranes achieved better salt rejection than non-coated membranes. Under the 

circumstances of very high salt concentrations, non-coated SWC3+ membranes experienced significant 

decline in salt rejection, while coated SWC3+ membranes maintained salt rejection values above 99.5%. 

No noticeable flux decrease was observed on both coated and non-coated SWRO membranes for this 

pilot. However, FTIR spectra of the used membrane samples indicated a thin fouling layer on both coated 

and non-coated RO surfaces, while coated sample had less foulant deposition than the non-coated 

membrane sample. Therefore, in conclusion, dopamine coated SWC3+ outperformed the non-coated 

membrane with higher salt retention and lower fouling potential and higher flux for UF stage.

Coated spiral UF membranes achieved sustainable performance in treating this produced water, 

although they required chemical enhanced cleaning operations a few times a day to maintain high (35 

GFD) flux. Simple acid and base enhanced hot water cleaning can efficiently recover the flux, while full 

standard CIP was occasionally needed. Foulants found on the membrane surfaces were mainly organic in 

nature. Considering the challenging raw water quality in this pilot test, this type of spiral UF membrane 

still demonstrated a relatively more easily cleaned surface property following application of a 

PDOPA/PEG coating.

4.2 Key Notes from the Pilot

(1) It is obvious that hollow fiber UF is a more efficient and easily maintained module than spiral 

wound configuration. Chemically enhanced HWC was only required every 12 hours for hollow fiber 

elements, while the cleaning frequency for spiral elements was every 4-6 hours. Routinely built-in 

backflush on the hollow fiber system maintained the flux in a more stable range than that of the spiral 

elements. 
38

(2) The high salinity of this produced water hampered the productivity of the RO procedure. With 

feed TDS value at 25,000 mg/L, 580 PSI feed pressure can achieve 50% recovery on water production. 

When feed TDS went above 50,000 mg/L, 800 PSI feed pressure can only achieve 15% recovery. In this 

case, one stage RO process becomes very inefficient and highly energy consumable in treating shale 

produced water. 

Nanofiltration (NF) membranes are widely used in brackish and seawater desalination. NF offers 

many advantages over RO in cases of high salinity feed, such as high flux with lower operating pressure, 

high retention of multivalent anions (Ca, Mg, Ba, Sr) and organic matter. Two-stage NF or using NF as 

RO pretreatment are frequently practiced in seawater desalination. Overall seawater TDS can be cut by 

57 %, while 93% of total hardness can be removed by one time pass of NF. The permeate obtained with 

NF would be much easier to treat by RO or a second stage of NF. It was reported that this integrated 

desalination system reduces the seawater desalination cost by 20-30% compared to conventional SWRO 

[22, 23] and 70% water recovery is possible from NF pretreated SWRO [24]. 

Recognizing the enhanced salt retention property of dopamine coated RO membranes in this 

pilot, we expect this coating technology would boost NF membrane salt rejection by a considerable

amount while the high permeation of the nanofiltration process could still be retained while 

simultaneously providing better anti-fouling characteristics for sustained, long term operation. 

Two stages of dopamine coated nanofiltration or one stage of coated NF followed by another 

stage coated SWRO is highly recommended to optimize the membrane filtration productivity for shale 

produced water treatment.

4.3 Future Work Recommendation

Based on this short-term pilot study, future pilot studies and research should be considered:

(1) A relatively long-term pilot should be considered for further evaluating the UF/RO dopamine 

coating technology in the application of shale produced water treatment. A twelve months run time 

is recommended for this long-term pilot using a two stage NF + NF or NF + RO system. In 

addition, for optimum operation for pre-treatment, we recommend testing of both MF and UF prior 

to RO using outside-in membranes. RO membrane fouling is expected to occur during longer tests,

and CIP recovery will provide more membrane cleaning efficacy for coating technology. Optimal 

membrane cleaning and system maintenance methods will also be resolved in such a pilot test. The 

cost of reclaimed water production from shale produced water can be accurately estimated from 

such pilot.

(2) Water quality needs to be routinely investigated for the next stage pilot test. Parameters like TOC 

and TPH should be more frequently monitored during the entire pilot test.

(3) Membrane filtration systems should be designed with the objectives set forth in bullet #1. We 

recommend that the future pilot test use hollow fiber ultrafiltration and microfiltration membranes,

and add at least one stage of nanofiltration membranes as the pretreatment for reverse osmosis 

process. This change can maximize the productivity of the permeate to the maximum extent 

possible. Meanwhile, maintenance work will be significantly reduced. With the help of this 
39

innovative anti-fouling coating technology, UF and multi-stage NF/RO or NF/NF membrane 

filtration, we expect to enable cost effective treatment of produced water to solve the water reuse 

problem in shale produced water.

5. Acknowledgements

We would like to acknowledge support from Devon Energy for providing access to their facilities 

and preprocessed water at their site near Fort Worth Dallas. We would also like to acknowledge the 

operations staff of Fountain Quail and Anthony Smith of Devon Energy who coordinates activities at the 

Maggie Spain Water Reuse Facility for their assistance in providing us a feed stream that was pre-treated 

via coagulation and clarification as well as pH adjusted. This work was supported by funding from The 

Research Partnership to Secure Energy for America (RPSEA) under Contract 08122-05. 

6. References

1. U.S. Geological Survey, Assessment of Undiscovered Oil and Gas Resources of the Bend Arch-Fort 

Worth Basin Province of North-Central Texas and Southwestern Oklahoma, 2003, USGS Fact Sheet 

(2004) 2004-3022. See:http://pubs.usgs.gov/fs/2004/3022/fs-2004-3022.pdf .

2. G. Clouser, The Booming Barnett Shale, Shale Gas, (2006) 8-12, January.

3. P. Galusky, Fort Worth basin/Barnett Shale Natural Gas Play: An Assessment of Present and 

Projected Fresh Water Use, Gas Technology Institute, (2007). Available at www.texerra.com.

4. Review of the Desalination and Water Purification Technology Roadmap, The National Academies 

Press, Washington, D.C., 2005.

5. R.F. Service, Desalination freshens up, Science, 313 (2006) 1088-1090.

6. T. Hayes, D. Arthur, Overview of Emerging Produced Water Treatment Technologies, International 

Petroleum Environmental Conference, (2004) October 12 IPEC Proceedings.

7. G. Doran, L.Y.C. Leong, Developing a Cost Effective Environmental Solution for Produced Water 

and Creating a “New” Water Resource, U.S. Department of Energy Report, (2000) DOE/MT/95008-

4.

8. T. Hayes, The Electrodialysis Alternative for Produced Water Management, GasTIPS, (2004) 

10(3):15-20. 

9. T. Hayes, Barnett and Appalachian Shale water Management Reuse Technologies, GTI Proposal 

#70100. 

10. H. Lee et al, Mussel-Inspired Surface Chemistry for Multifunctional Coatings, Science 318 (2007) 

426-430. 

11. D.R. Dreyer, D.J. Miller, B.D. Freeman, D.R. Paul, and C.W. Bielawski, Elucidating the Molecular 

Structure of Poly(dopamine), in preparation.

12. J.T. Arena, B. McCloskey, B.D. Freeman, J.R. McCutcheon, Surface Modification of Thin Film 

Composite Membrane Support Layers with Polydopamine: Enabling Use of Reverse Osmosis 

Membranes in Pressure Retarded Osmosis, Journal of Membrane Science, in press.

13. B. D. McCloskey, H.B. Park, H. Ju, B.W. Rowe, D.J. Miller, B.J. Chun, K. Kin, and B.D. Freeman, 

Influence of Polydopamine Deposition Conditions on Pure Water Flux and Foulant Adhesion 

Resistance of Reverse Osmosis, Ultrafiltration, and Microfiltration Membranes, Polymer, 51, (2010), 

3472-3485.
40

14. What Is Membrane Performance Normalizations? Hydranautics technical supporting material. 

Available at http://www.membranes.com/docs/trc/normaliz.pdf.

15. KMS HFM PVDF Food & Dairy UF Elements Specification. Koch Membrane Systems technical 

supporting material. Available at http://www.kochmembrane.com/support_uf_lit.html.

16. H. Choi et al., Effect of Permeate Flux and Tangential Flow on Membrane Fouling for Wastewater 

Treatment, Separation and Purification Technology 45 (2005) 68-78.

17. C. Tang, Y. Kwon, J. Leckie, Probing the Nano- and Micro-scales of Reverse Osmosis Membranes –

A Comprehensive Characterization of Physiochemical Properties of Non-coated and Coated 

Membranes by XPS, TEM, ATR-RTIR, and Streaming Potential Measurements, Journal of 

Membrane Science, 287(1) (2007), 146-156.

18. D. L. Pavia, G.M. Lampman, G. S. Kriz, Introduciton to Spectroscopy: A Guide for Students of 

Organic Chemistry, W. B. Saunders Company, Philadelphia, 1979.

19. K. Khatib et al, Physico-chemical Study of Fouling Mechanisms of Ultrafiltration Memebrane on 

Biwa Lake (Japan), Journal of Membrane Science, 130 (1997), 53-62.

20. F. J. Stevenson, Humus Chemistry: Genesis, Composition, Reactions, 2nd ed., Wiley, New York, 

1994.

21. A. Rao et al., Structure-performance Correlation of Polyamide Thin Film Composite Membranes: 

Effect of Coating Conditions on Film Formation, Journal of Membrane Science, 211(1), (2003), 13-

24.

22. C. K. Diawara, Nanofiltration Process Efficiency in Water Desalination, Separation and Purification 

Reviews, 37:3, (2008), 302-324.

23. Seawater Desalination - The Long Beach Method. Available at: 

http://www.lbwater.org/pdf/desal_lbmethod.pdf.

24. A. M. Hassan et al., A New Approach to Membrane and Thermal Seawater Desalination Processes 

Using Nanofiltration Membranes (Part-1), Desalination and Water Resuse Quarterly, May/June 

Issure, (1998), 8/1, 53-59.