Methods of making purified water from the Fischer-Tropsch process
11739014 · 2023-08-29
Assignee
Inventors
Cpc classification
C02F2201/46115
CHEMISTRY; METALLURGY
C02F2103/365
CHEMISTRY; METALLURGY
Y02E50/30
GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
C02F1/20
CHEMISTRY; METALLURGY
Y02W10/10
GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
C02F2305/06
CHEMISTRY; METALLURGY
International classification
Abstract
The Fischer-Tropsch (FT) process creates significant amounts of water. The invention provides an aqueous composition comprising specified amounts of dissolved and suspended solids, low chemical oxygen demand and low chlorine demand, pH in the range of 6.5 to 8.0, where the aqueous composition comprises organic carbon derived from fossil sources.
Claims
1. An aqueous composition, comprising Total Dissolved Solids (TDS) of 100 to 300 mg/l; dissolved salts wherein 90 mass % or more of the dissolved solids salts are sodium bicarbonate or potassium bicarbonate; Total Suspended Solids (TSS) of less than 5 mg/l; Total Organic Carbon (TOC) of 1.0 to 10 mg/l; Chemical Oxygen Demand (COD) of less than 50 mg/l; chlorine demand of less than 5 mg/l; pH in the range of 6.5 to 8.0; Hardness of less than 50 mg/l as CaCO.sub.3; and wherein the organic carbon in the aqueous composition is derived from fossil sources.
2. The aqueous composition of claim 1 comprising 150 to 300 mg/l TDS.
3. The aqueous composition of claim 1 comprising a TOC of 2.0 to 8 mg/l.
4. The aqueous composition of claim 1 wherein the organic carbon consists essentially of C1 to C10 monohydroxy-alcohols, carboxylic acids and polysaccharides.
5. The aqueous composition of claim 1 wherein the organic carbon comprises C1 to C10 monohydroxy-alcohols or carboxylic acids.
6. The aqueous composition of claim 1 comprising a COD of 1.0 to 50 mg/l; or 5.0 to 50 mg/l; or 2.0 to 40 mg/l.
7. The aqueous composition of claim 1 having a hardness of 1 to 50 mg/l, or 5 to 50 mg/l, or 2 to 40 mg/l as CaCO.sub.3.
8. The aqueous composition of any of claims 1, 2, and 3 wherein the organic carbon comprises C1 to C10 monohydroxy-alcohols.
Description
BRIEF DESCRIPTION OF THE DRAWINGS
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DETAILED DESCRIPTION OF THE INVENTION
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(8) A flow diagram of an overall process (100) is illustrated in
(9) In some preferred aspects of the invention, some or all of the water created in the FT process is subjected to a stripping operation. In some preferred aspects, the stripper pressure is slightly above atmospheric pressure and the temperature of the mixture at any point in the column will be at the mixture bubble point. In some preferred aspects, the stripping can be done by flowing the FT water down a column with packing or trays, with the stripping fluid (e.g. injection of live (i.e., pressurized) steam, nitrogen, air, or other available gases or vapour) in counter-current contact. Alternatively, heat may be supplied to the stripper by reboiling a portion of stripped water. Alternatively, or in addition, the water could be distilled; however, stripping is preferred. The stripping may be done under vacuum or slightly above atmospheric pressure (for example, 0.1-10 atm). The temperature will be below the boiling temperature of the FT water. The mass ratio of stripping medium to FT water may be 0.001 to 0.5, more preferably 0.01 to 0.2.
(10) Water created in an FT process conducted at contact times of 5 seconds or less, more preferably 2 s or less or is or less and/or short diffusion distances (e.g. FT catalyst coating thickness of 100 μm or less, or FT catalyst particle size of 1,000 microns or less, or less than 500 microns or less) can be superior to water created by conventional FT or many other industrial waste water compositions. Advantages of the created water obtained as described herein may include one or more of the following features: very low concentration of aromatics (e.g., 10 ppm or less); low aldehyde concentration, and wherein the carbon present in the water is nearly exclusively (e.g., at least 90% by mass, or at least 95% by mass, or at least 98% by mass, or at least 99% by mass) in the form of biodegradable acids (e.g., formic, acetic, propionic, n-butyric, n-valeric, and caproic), or alcohols (e.g., methanol, ethanol, propanol, butanol, decanol).
(11) FT processes that are conducted in microchannels comprising an FT catalyst and/or at short contact times with an FT catalyst are especially desirable since such processes result in a superior mix of components dissolved in the FT produced water as compared to conventional FT processes. Further, the mixture of oxygenated species in these short contact time FT processes enable simple water treatment using this invention without requiring the need for stripping columns required for more difficult separations. The FT produced water in this invention may be processed in a simple stripper which does not require either (1) a side draw of vapor or liquid from the stripper or (2) liquid-liquid phase separation of the condensed overhead stream. The stripper may be operated such that the composition of the distillate product is the same as the composition of the reflux. For purposes of the present invention, a microchannel is defined as a channel having at least one internal dimension of 10 mm or less; in some preferred embodiments 5 mm or less. In preferred embodiments, the FT reaction is conducted through a planar array of microchannels that are adjacent to a planar array of coolant channels. Short contact times are preferably less than 5 second, more preferably less than 500 msec, and in some embodiments in the range of 150 to 500 ms.
(12) While this process is useful for microchannel reactors, the methods provide significant advantages for any FT process, whether using a conventional reactor or not. There is a particular advantage for small-scale facilities that produce 15,000 barrels per day (BPD) or less, preferably 5000 BPD or less, of FT liquids and solids. The reason for this advantage is that at the small scale of the facility, scaling down prior art waste water treatment processes is difficult and costly. There is a need to have a very simple waste water treatment process that can be built using modular construction. The stripper and the membrane reactor and associated equipment can be built on modular structures in separate facilities and transported to the site by truck, rail, or shipping. These systems can be used to avoid construction associated with conventional waste water treatment processes, such as waste water ponds using concrete structures or in-ground retention ponds with liners. For these reasons, the invention is also useful for treating FT water for grass-roots facilities or facilities in remote locations where it is difficult to integrate the waste water treatment with existing facilities, such as an existing refining waste water treatment facility. Thus, the invention includes modular components for the FT process and/or water treatment including the MBR and other components. The invention also includes a kit for transporting the modular components. With this in mind, the invention is useful for any FT reactor type, whether conventional (slurry or fixed bed) or a new technology such as compact, structured reactors, including microchannel reactors.
(13) The Fischer-Tropsch Process
(14) Examples of Fischer-Tropsch processes suitable for use in the present invention are described in U.S. Pat. Nos. 9,023,900, 7,935,734 US Published Patent Application No. 2014/0045954 and WO2012107718 which are incorporated herein by reference. The following are some non-limiting descriptions of some preferred embodiments of the FT process that can be used for creating water in conjunction with the present invention.
(15) Suitable apparatus for conducting the FT process is known in the prior art. Preferred apparatus are microchannel reactors. A microchannel reactor may be made of any material that provides sufficient strength, dimensional stability and heat transfer characteristics to permit operation of the desired process. These materials may include aluminum; titanium; nickel; platinum; rhodium; copper; chromium; alloys of any of the foregoing metals; brass; steel (e.g., stainless steel); quartz; silicon; or a combination of two or more thereof. Each microchannel reactor may be constructed of stainless steel with one or more copper or aluminum waveforms being used for forming the channels. In preferred embodiments, the FT reactor is not a fluidized bed reactor.
(16) The FT reactor may comprise a plurality of plates or sheets in a stack defining a plurality of Fischer-Tropsch process layers and a plurality of heat exchange layers, each plate or sheet having a peripheral edge, the peripheral edge of each plate or shim being welded to the peripheral edge of the next adjacent plate or shim to provide a perimeter seal for the stack. Some preferred construction techniques are shown in U.S. application Ser. No. 13/275,727, filed Oct. 18, 2011, which is incorporated herein by reference.
(17) The FT reactor may be constructed using waveforms in the form of corrugated inserts. These corrugated sheets may have corrugations with right-angles and may have rounded edges rather than sharp edges. These inserts may be sandwiched between opposing planar sheets or shims. In this manner the channels may be defined on three sides by the corrugated insert and on the fourth side by one of the planar sheets. The process microchannels as well as the heat exchange channels may be formed in this manner. FT reactors made using waveforms are disclosed in U.S. Pat. No. 8,720,725, which is incorporated herein by reference.
(18) The FT reactor may comprise at least one process channel in thermal contact with a heat exchanger, the catalyst being in the process channel. The reactor may comprise a plurality of process channels and a plurality of heat exchange channels, the catalyst being in the process channels.
(19) The catalyst is preferably in the form of particulate solids. These particulates can be packed into parallel arrays of small channels (typically having a width and/or height dimension of 1 cm or less, preferably 1 mm to 1.0 cm, and any length, for example lengths of 50 cm or 1 m or greater) that are interleaved with parallel arrays of heat exchange channels. Alternatively, the catalyst may be coated on interior walls of the process channels or grown on interior walls of the process channels. The catalyst may be supported on a support having a flow-by configuration, a flow-through configuration, or a serpentine configuration. The catalyst may be supported on a support having the configuration of a foam, felt, wad, fin, or a combination of two or more thereof. The catalyst may comprise a coating on a monolith, including monoliths that may be separately inserted or removed from the reactor.
(20) In one preferred process for conducting a Fischer-Tropsch reaction, a reactant mixture in a reactor flows in contact with a catalyst to form a product comprising at least one higher molecular weight hydrocarbon product. Preferably, the catalyst is derived from a catalyst precursor comprising cobalt, optionally a promoter such as Pd, Pt, Rh, Ru, Re, Ir, Au, Ag and/or Os, and a surface modified support, wherein the surface of the support is modified by being treated with titania, zirconia, magnesia, chromia, alumina, silica, or a mixture of two or more thereof. FT processes with short contact times are enabled by high cobalt catalyst loadings, such as catalyst with greater than 20 mass %, more preferably greater than 25%, 35%, or greater than 50 mass % cobalt loading. The product further comprises a tail gas, and at least part of the tail gas can be separated from the higher molecular weight hydrocarbon product and combined with fresh synthesis gas to form a reactant mixture, the volumetric ratio of the fresh synthesis gas to the tail gas in the reactant mixture being in the range from about 1:1 to about 10:1, or from about 1:1 to about 8:1, or from about 1:1 to about 6:1, or from about 1:1 to about 4:1, or from about 3:2 to about 7:3, or about 2:1; the reactant mixture comprising H.sub.2 and CO, the mole ratio of H.sub.2 to CO in the reactant mixture based on the concentration of CO in the fresh synthesis gas being in the range from about 1.4:1 to about 2:1 or from about 1.5:1 to about 2.1:1, or from about 1.6:1 to about 2:1, or from about 1.7:1 to 1.9:1.
(21) In some preferred embodiments, all water produced in the FT process is collected in a separator, vessel, or tank, and the full flow is subjected to stripping, prior to subsequent use and/or treatment in an MBR.
(22) The MBR reactor should not be configured with long stretches of plug flow. This is because, in plug flow, the pH will rise as the acids are consumed by the microorganisms. Instead, mixing should be permitted such that pH remains similar throughout the MBR's volume. In one embodiment, a MBR can be run with a F:M of 0.05 to 0.5, a sludge age of 5 to 50 days, a pH in the range of 5 to 9; a temperature in the range of 20 to 40 C, and a conductivity of 50 to 1500 μS/cm.
(23) Except for the specified conditions mentioned herein, conditions in the MBR are conventional. For example, the membranes can be removed and washed with dilute sodium hypochlorite, as is conventional in the art. Microorganisms can be sourced from any municipal or industrial activated sludge plant. The bacteria will adapt over time to the feed source, and produce a simple, readily biodegradable waste which is suitable for most heterotrophic and indeed autotrophic bacteria. There is no need for a special source of microbes. Preferably, the MBR should be run at steady state conditions meaning constant feed rate, constant sludge age, constant pH and so on; avoiding large swings in operating conditions.
(24) The overall process is typically considered to be continuous, but in typical commercial operation, plural MBRs will operate in parallel, and each MBR will be brought down from time-to-time for cleaning of the membranes. During this period, the total flow will be accommodated within the online MBR's, although at slightly higher hydraulic load.
(25) Since the FT produced water lacks nutrients, nutrients will need to be added for the microorganisms. In a preferred embodiment, nitrate, potassium, calcium, magnesium, manganese, sodium, iron, copper, zinc, molybdenum, nickel, and cobalt are added. For example, one preferred nutrient mix comprises ammonium nitrate, potassium nitrate, calcium nitrate, magnesium nitrate, manganese chloride, iron chloride, copper chloride, zinc chloride, nickel chloride, cobalt chloride and sodium molybdate.
EXAMPLES
(26) Described below is project aimed to establish the efficiency of treating FT produced waterin an MBR. Two MBR pilot plants were employed, one fed with an analogue and the other with FT produced water which was stripped of the volatile fraction. Both treatability (in terms of COD removal) and performance are assessed based on the COD removal and sustainable flux under different operating conditions.
(27) 1 Materials and Methods
(28) 1.1 FT Produced Water
(29) Around 2000 L of FT produced water (Table 1-1) was shipped from pilot GTL facility in Ohio, US. Samples were collected from separators and contained COD concentration of 25-30 g/l and pH 3.3. About 75% of the COD is made up by methanol and ethanol.
(30) The absence of minerals in the FT produced water, including nitrogen and phosphorus (N & P) demanded dosing with micro-nutrients to sustain biomass growth (Tehobanoglous et al, 2004). For the smaller MBR pilot plant, a nutrient mix was prepared (Table 1-2) and blended the feed water. For the larger plant, a commercial nutrient mix was used and dosed into the dilution water. Since there was no natural pH buffering, a small dose of caustic soda was added to the feed to provide a sodium bicarbonate buffer in the bioreactor.
(31) TABLE-US-00006 TABLE 1-1 FT produced water Composition Composition mg/L COD 30100 PH 3.29 TOC 6600 DI-Ethyl- 181 Ketone Alcohols c1-c10 Butanol 446 Decanol 5.4 Ethanol 2450 Heptanol 28 Hexanol 94.2 Methanol 7780 Nonanol 5.2 Octanol 10.2 Pentanol 251 Propanol 690 Organic acids Acetic acid 261 Propionic acid 20.4 Formic Acid 108 n-Butyric acid 22.8 n-Valeric acid 23.6 Caproic acid 19.4 Aldehydes Acetaldehyde <10 Butyraldehyde <10 Glutaraldehyde <10 Glutaraldehyde <10 Isobutyraldehyde <10 Propionaldehyde <10 Formaldehyde <10 VOC Composition ug/L 2-Butanone 8620 Acetone 52300 DEK 181 M,P-Xylene 8.3 Styrene 4.9 Strene 2.4 o-Xylene 5.4
(32) TABLE-US-00007 TABLE 1-2 SMBRs Nutrient Composition (mg/g dry solid) Composition mg Composition mg Composition mg N 104 Mn 0.1 K 12 Mo 0.0048 Zn 0.2 S 25 Cu 0.024 Fe 2.4 Ca 12 Co 0.00048 P 21 Ni 0.001 Mg 8
In some preferred embodiments, the nutrient composition can be defined as having each element within 50% to 200% of the concentrations shown in Table 1-2.
Examples of FT produced water made using a short contact time FT process are shown in the Tables below:
(33) TABLE-US-00008 TABLE 3 Example FT reaction water composition Component Units Normal load Peak load Acetone mg L.sup.−1 28 53 Methyl ethyl ketone mg L.sup.−1 5.0 8.7 Diethyl ketone mg L.sup.−1 0.0 0.2 Benzene μg L.sup.−1 1.3 0.0 Toluene μg L.sup.−1 8.8 2.4 Xylene μg L.sup.−1 4.6 14 Styrene μg L.sup.−1 0.0 5.0 Acetaldehyde mg L.sup.−1 10 0.0 Formic acid mg L.sup.−1 78 108 Acetic acid mg L.sup.−1 478 261 Propionic acid mg L.sup.−1 86 20 Butyric acid mg L.sup.−1 76 23 Valeric acid mg L.sup.−1 76 24 Hexanoic acid mg L.sup.−1 54 20 Methanol mg L.sup.−1 4038 7779 Ethanol mg L.sup.−1 1935 2450 Propanol mg L.sup.−1 460 690 n-Butanol mg L.sup.−1 232 446 n-Pentanol mg L.sup.−1 163 251 n-Hexanol mg L.sup.−1 80 94 n-Heptanol mg L.sup.−1 40 28 n-Octanol mg L.sup.−1 9.2 10 n-Nonanol mg L.sup.−1 5.8 5.0 n-Decanol mg L.sup.−1 0.0 5.4 TPH mg L.sup.−1 7.1 7.1
(34) TABLE-US-00009 TABLE 4 Example FT reaction water composition Component Units Normal load Peak load Acetone mg L.sup.−1 28 53 Methyl ethyl ketone mg L.sup.−1 5.0 8.7 Diethyl ketone mg L.sup.−1 0.0 0.2 Benzene μg L.sup.−1 0.0 1.3 Toluene μg L.sup.−1 2.4 8.8 Xylene μg L.sup.−1 4.6 14 Styrene μg L.sup.−1 0.0 5.0 Acetaldehyde mg L.sup.−1 0.0 10 Formic acid mg L.sup.−1 101 129 Acetic acid mg L.sup.−1 365 479 Propionic acid mg L.sup.−1 44 79 Butyric acid mg L.sup.−1 39 64 Valeric acid mg L.sup.−1 37 62 Hexanoic acid mg L.sup.−1 26 41 Methanol mg L.sup.−1 8900 11782 Ethanol mg L.sup.−1 3088 3821 Propanol mg L.sup.−1 954 1411 n-Butanol mg L.sup.−1 593 803 n-Pentanol mg L.sup.−1 336 456 n-Hexanol mg L.sup.−1 118 154 n-Heptanol mg L.sup.−1 61 88 n-Octanol mg L.sup.−1 20 31 n-Nonanol mg L.sup.−1 9.4 15 n-Decanol mg L.sup.−1 7.7 15 TPH mg L.sup.−1 7.1 7.1
Water compositions may be determined by GC/mass spectrometry or other appropriate techniques. Care should be taken to avoid vaporization during sampling. It is believed that the values in Table 4 are more accurate than the values in Table 3.
1.2 Small MBR (SMBR)
1.2.1 Bench ScaleSetup
The small MBR (
(35) The raw FT produced water was steam stripped prior to biotreatment to remove the bulk of the volatile organic carbon (VOC) and so reduce the organic load for aerobic degradation. At full-scale this would be conducted using a packed tower stripper. At bench scale, the FT produced water was boiled continuously at 100° C. for an hour to remove the alcohols, to reduce the COD concentration from 25-30,000 mg/L to around 4000 mg/L. It was diluted with deionised water to achieve the desired COD level of 1000-5500 mg/L depending on the desired experimental conditions. A COD of 1,000 mg/l was achievable with boiling alone, but the loss of water by evaporation is high due to the extended boiling period. In a full scale steam stripping column, water loss is avoided by condensing the stripper overheads. This was represented by simply diluting the partially stripped water with deionised water (DI). Nutrient (Table 1-2) was dosed separately to avoid bacterial growth in the feedstock.
(36) The SMBR was seeded with 4 L activated sludge from a bioreactor (an industrial SBR treating bottling plant wastewater) that had been pre-acclimatised to FT produced water. During start-up the MLSS concentration was increased to approximately 10 g/L. The operating conditions were subsequently adjusted (Table 1-5) to sustain different MLSS concentrations and F:M ratios of 0.19-0.3 d.sup.−1. 0.3 to 0.19 from this project have little impact on the effluent quality and the operation range from 0.17 to 0.32 have shown stability in the MBR performance
(37) TABLE-US-00010 TABLE 1-5 SMBR Operating Parameters Feed-COD MLSS HRT SRT FLUX mg/L mg/L F:M Hours Days LMH 2500 12000 0.1 40 50 1 5500 12000 0.3 35 50 1 4000 15000 0.2 30 50 1 3000 15000 0.25 12 35 3.5 2000 15000 0.3 12 35 3.5 2500 17000 0.3 12 35 3.5 1500 11000 0.25 12 20 3.5 1200 8500 0.25 12 20 3.5
1.3 Large MBR (LMBR)
1.3.1 Pilot Plant Set-Up
(38) The LMBR had a nominal capacity of 150 L and was fitted with 5 flat sheet MF panels identical to that of the SMBR, providing a total area of 0.5 m.sup.2. The MBR temperature was controlled at ˜30° C. using a heating jacket. A coarse bubble membrane tube air diffuser delivered 100 L/min to sustain a minimum DO of 2 mg/L, as monitored manually. The diffuser was located at the base of the membrane cassette, to provide membrane air scouring of 12 Nm.sup.3/(h.Math.m.sup.2) at 1001/min as well as provide oxygen for the process. Peristaltic pumps (Watson Marlow 101U/R) were used to pump the feed and draw permeate as with the SMBR. A GAC polishing step was incorporated based on 750 g of Norit GAC 1240 W (Steam activation of coal).
(39) The SMBR feedwater comprised a combination of unstripped FT produced water blended with acetic acid, the ratio of FT produced water to Acetic acid was 52:1, recycled permeate and a solution of commercial botanic nutrient (Miracle-Gro), diluted with 200 L of de-chlorinated potable water. For the first 2 months, diluted unstripped FT produced water alone was used as the feed source. After 2 months the feed was supplemented by adding 9 kg of acetic acid and 2.5 kg of propionic acid, both reagent grade, and topping up with tap water. Unstripped FT produced water with and without spiking was diluted to obtain the required COD test concentration. Feed and nutrient were dosed from 2 different tanks, and recycled permeate dosed directly back to the MBR, to avoid cross contamination causing bacterial growth. The LMBR was seeded with activated sludge from a municipal WWTP with MLSS of approximately 7 g/L, and the MLSS subsequently gradually increased to 9 g/L before adjusting further according to the experimental programme (Table 1-6).
(40) TABLE-US-00011 TABLE 1-6 LMBR Operating Parameters Feed- COD MLSS F:M HRT SRT FLUX 2000 9000 0.2 18 NC 17 2000 12500 0.2 18 NC 17 2000 9000 0.2 25 36 11 2000 10000 0.2 25 30 11
1.4 Membrane Cleaning
(41) Operation was sustained without routine chemical cleaning in place (CIP) and fouling behaviour observed with reference to the pressure, monitored using an analogue vacuum gauge on the pump suction line. Recovery cleaning was performed when the pressure reached 0.3 bar. Membranes were then removed and washed at low pressure with mains water before applying mechanical cleaning with a sponge and then soaking in 1000 mg/L hypochlorite for 30 mins and then rinsing with deionised water.
(42) 1.5 Analytical Methods
(43) The Chemical Oxygen Demand (COD) was tested using Merck's Cell Test and Spectroquant Photometer NOVA 60 according to Standard Methods (APHA, 2005). Standard APHA methods were also used to estimate mixed liquor suspended solids (MLSS), mixed liquor volatile suspended solids (MLVSS), total dissolved solids (TDS), Capillary Suction Time (CST) and chlorine demand. The CST readings were obtained using Triton 2000 series CST filterability tester and Triton CST (7×9 cm) filter paper. Chlorine demand was measured using Hach Colorimeter Filter Photometer in combination with Hach DPD Total Chlorine Reagent Powder Pillows, 0.02-2.00 mg/L range.
(44) The GAC isotherm determination employed GAC particles fractioned to 32-106 μm in size at masses of 0.1, 0.5, and 1 g in a 120 mL volume of SMBR permeate. The suspensions were shaken for 6 hours and the solution sampled for residual COD.
(45) As is well known, pH can be measured using conventional metering apparatus and techniques.
(46) 3 Results and Discussion
(47) 3.1 Water Quality Vs. Retention Time
(48) In the initial phase of SMBR operation the MLSS was allowed to increase to a maximum of 18 g/L at an HRT of 32-42 h. During this period a steady and gradual improvement of COD removal from 97% to 99% was observed when operation was stable (feed COD 3.1-3.9 g/L) and the MLSS between 14 and 16 g/L. The SRT was decreased from 33 days to 20 days to reduce the MLSS to around 8.8 g/L. This resulted in deterioration in COD removal in the short term (22 days) when operating conditions were being changed. However, on returning to steady-state operation of 20d SRT, 12 h HRT and 1,000-1,200 mg/L COD feed the COD removal increased to >99%.
(49) The LMBR MLSS was increased from 7 g/L to 15 g/L in first month of operation, with no sludge wastage up to 12.5 g/L MLSS at an HRT of 17.5 h. During this initial period the COD removal rate was stable between 97%-98% (
(50) Overall COD removal of >99% was demonstrated for the stripped wastewater for steady-state conditions. This corresponded to a COD concentration as low as 5 mg/L in the MBR-treated wastewater. The recorded COD removal data from this study were in line with those from the two other reported FT produced water treatment studies. They are considerably greater than data reported for most of the bench-scale studies and full-scale plant operation (Table 3-1), reflecting the benign nature of FT wastewater compared with other O&G effluents which are generally much more challenging (Table 0-2 and Table 3-1). The COD removal rate for the current study meets the World Bank Group's industry guidelines of 150 mg/L maximum COD (WBG, 2007) for petroleum industry effluents.
(51) Operation of the MBRs was generally disrupted by malfunctions which would not be expected to arise in a full-scale plant. Foremost amongst these was the clogging of the feedwater tubes of the SMBR in particular, causing significant fluctuation in the biotank sludge volume and MLSS concentration. Clogging was caused by both the precipitation of ferric oxide—an unanticipated contaminant in the feedwater—and the formation of biofilms (including algal growth) associated with the nutrient feed dosing. An unusual facet of the biotreatment of the FT wastewater is the pH trend, where the treated wastewater is more alkaline than the feed. Whereas the feed wastewater pH was between 4 and 5, the treated wastewater pH at steady-state (maximum COD removal) was around 8 for the LMBR and 7.5 for the SMBR. This is because the acidity generated by the carboxylic acids in the feed is removed once the acids are degraded.
(52) TABLE-US-00012 TABLE 3-1 Case studies Comparison on Petroleum Industry's MBR Parameter and Performance. COD Country/ Effluent Membrane Temp, Flux, SRT, HRT, MLSS, Feed, COD Company Region Scale Source Type ° C. LMH Days Hrs g/L g/L Removal Reference Sinopec China Full Oil Refining, iHF 20- 8.9 30- 15 3- 225 80% Judd, Guangzhou Ethylene Process 40 90 8 (min) 2014 Petrochemical China Full Petrochemical iHF 15- 16 15- 0.33 6 2500 98% plant, Sichuan 30 30 (Max) Formosa, Yunlin Taiwan Full Oil Refining, iHF 20- 20- NA NA 3.5 1000 95% Petrochemical 30 30 Syndial, Italy Full Ethylene/ iHF NA 16- 15- NA 6 280 58% Porto Marghera PVC Process 19.8 30 (min) Sangachal Azerbaijian Pilot Offshore Oil iHF 15- 13- NA NA 20 4000- 97% Rees et Terminal, Baku Reserviors 27 19.3 50000 (min) al., 2009 TPAO Basin, Turkey Bench PW iHF 20 10 30- 2.7.sup.b 2- 1000- 67- Kose et Trakya “inf” 16 3000 83.sup.a al., 2012 Petrochemical Singapore Bench Petrochemical iFS 26 10- 25 13- 8.6- 720- 85- Qin et plant 12.5 16 9.6 1590 95.sup.d al., 2007 Petrobras oil Brazil Bench Refinery iHF ~25 15- “inf” 10.0 2.1- 400- 41- Viero et refinery 17.5 10.4 1050 67 al., 2008 Queensland Energy Australia Bench Oil Shale Retort iFS 25 0.6 N/A 168- 14- 10,000- 75- Lea et resources Sour Water 504 20 30,000 80% al 2013 Sasol Technology, South Pilot GTL iFS 42 17 35 8 7.8 1800 96% Young et Secunda Africa al., 2006
3.2 Sludge Quality
(53) SVI and CST were measured for the steady state mixed liquor. The SVI was not measurable, with no visible settling over a 2 hour period for either MBR. The LMBR had a mean CST of 316, ranging between 276 and 372 s over the final 4 weeks of the study. The SMBR CST values varied significantly, progressively increasing from 109 s to 1064 s over the same period. These values are significantly higher than those OF 5-50 s reported for FT produced water (Molipane et al., 2006) but are similar to those reported by Wu et al (2009) for Municipal sludge processed in the simultaneous sludge thickening and digestion reactor. The high CST and SVI values imply that thickening and dewatering of the sludge may be somewhat challenging (Fabregat et al, 2011).
(54) 3.3 Membrane Performance
(55) Practical constraints of the study meant that the SMBR flux was very low at ˜3.5 LMH, such that no fouling was observed for this MBR throughout the study. The LMBR was, however, designed to permit higher fluxes under more representative operating conditions of 11-17 LMH, albeit with an extremely high membrane air scour rate of around 12 Nm.sup.3/(h.Math.m.sup.2) compared with 0.2-0.8 normally associated with full-scale industrial effluent treatment MBR plants, (Judd, 2014). This flux applied is in line with the values listed in Table 3-1.
(56) The LMBR membrane ran without cleaning for the first month, but then required cleaning after each 7-11 day period of operation when the pressure increased to the threshold maximum value of 0.3 bar. Inspection of the membrane showed that it had clogged (
(57) 3.4 Post-Treatment
(58) Post-treatment using the GAC contributed only 0.35-1.45% to the total COD removal. The adsorption isotherm measurements indicated that increasing the concentration of GAC from 0.5 to 1 mg/L at the feed concentration of 12 mg/L COD had no impact on the equilibrium COD concentration of 6.4 mg/L COD±10%. It was therefore concluded that GAC was not a viable option for polishing residual dissolved COD, probably reflecting the low molecular weight of the residual organic carbon.
(59) The chlorine demand of the effluent was determined as being 0.4-0.46 mg/L. This means that the water can be disinfected for use as cooling tower make up or other services without incurring excessive costs for chlorination.
DISCUSSION
(60) For a high COD loading of up to 5500 mg/L a COD removal of over 98% was consistently achieved for both the real and analogue FT produced water, somewhat higher than previously reported values for petroleum industry wastewaters generally. This reflects the highly biodegradable nature of this wastewater, in marked contrast with other oil and gas effluents. Significant fluctuations in the F:M ratio did not affect the COD removal rate under conditions of progressively increasing the MLSS concentration. A flux of up to 14 LMH was sustained under operating conditions of high but uneven membrane air scour applied despite an extremely high sludge CST value of 275-372—indicating an innately high sludge filtration resistance. The low residual COD concentration of 5 mg/L or less attainable under steady-state operating conditions means that the treated effluent is suitable for re-use as cooling water following moderate doses of chlorine (0.4 mg/L chlorine demand) without requiring further polishing. The use of a completely mixed reactor permitted partial neutralisation of the feed with caustic soda, reducing the TDS of the effluent and decreasing the blowdown (waste stream) from its reuse for evaporative cooling. MBRs would provide a lower-footprint process than the previously applied membrane polishing process (classical activated sludge followed by membrane filtration). Although the previous patent literature disclosed that pH in the MBR should be maintained between pH 5.5 to 9.5; surprisingly, we discovered that excellent results could be obtained with a relatively low pH of the water prior to addition to the MBR (i.e., the water feed to the MBR). And the use of a lower pH further improves the process by reducing or eliminating the amount of brine that must be disposed. It was also surprising that the pH increased in the MBR since in most industrial wastewater plants, the pH goes down through the treatment process.
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