A method for treating oil and gas field exploitation wastewater

Abstract

The application discloses a treatment method for oil and gas field exploitation wastewater, and relates to an improvement of a treatment method for oil and gas field exploitation wastewater, which comprises the following steps: 1, impurity removal; 2, special membrane oil removal; 3, desilication, heavy metal and sulfate removal; 4, degassing, removal of methane, methanol, inorganic ammonia and small molecule organic matters in the wastewater; 5, heterogeneous ozone catalytic oxidation; 6, membrane salt separation, membrane separation of monovalent salt and divalent salt; 7, evaporation crystallization; and 8, sodium chloride electrolysis. The method has the beneficial effect that the oil and gas field exploitation wastewater is maximally resourced. Only a small amount of solid waste is generated, four products, namely, sodium chloride, potassium chloride, calcium chloride / magnesium chloride (snow dissolving agent) and industrial raw oil, are obtained, and all the water is discharged up to the standard.

Description

Technical Field

[0001] This invention belongs to the field of environmental protection technology and relates to an improvement on a method for treating wastewater from oil and gas field extraction. Background Technology

[0002] Traditional processing methods:

[0003] 1) After treatment, the solution is reinjected into the oil extraction area and then injected underground.

[0004] 2) Resource utilization plan.

[0005] Disadvantages of traditional processing methods:

[0006] 1) Reinjection: Favorable reinjection areas are mostly located inside or around the gas reservoir, posing a potential threat to the deployment of new wells and the integrity of production well casings within the reservoir area; the reinjection distance to the production area is long, requiring passage through complex environments such as tunnels, rivers, and farmland, resulting in high investment costs for water diversion pipelines; the produced wastewater from the gas field is high in oil, salt, and mineralization, which is highly corrosive to reinjection pipelines and pumps, leading to high maintenance costs for water injection. Furthermore, the variation in water production in oil and gas fields is unpredictable, and the long design and construction cycles of water diversion pipelines cannot cope with changes in water volume.

[0007] 2) Resource Utilization Scheme: Current resource utilization schemes all require chemical addition in the pretreatment process. This significantly increases the salt content of the already high-oil, high-salt, and high-mineralization wastewater, raising treatment costs. The residual chemicals create numerous unresolved problems for the membrane and evaporation processes. Furthermore, the large amounts of solid waste and mixed salts generated during the treatment process cannot be digested. Summary of the Invention

[0008] The purpose of this invention is to achieve optimal treatment costs by eliminating the need for large amounts of reagents in the pretreatment of produced water from oil and gas fields for resource recovery. This is achieved solely through operational condition control and a proprietary process to separate and purify the components. Both the produced water and solids meet national / industry standards, with no hazardous waste or salt byproducts. This addresses the pain points of solid waste disposal.

[0009] The method of the present invention is:

[0010] 1. Impurity removal:

[0011] 1.1 The pH value of the gas field produced wastewater is adjusted to 2-3 using a high-speed mixer. The acid used to adjust the pH value comes from hydrochloric acid produced by the electrolysis of sodium chloride in the subsequent process. This reduces the solubility of oil in water, increases the solubility of inorganic salts in water, and fixes the upper limit of silicon solubility in water to ≤85mg / L.

[0012] 1.2 The oil-containing water is separated and concentrated using a tubular / plate filter. The separated clear water enters the next oil removal process. The separated reduced-volume slurry is pressed into a filter cake by a diaphragm filter press. The filter cake is washed and dried with the qualified water recovered from the evaporation and crystallization process and treated as solid waste. The washing water is returned to the filter for re-filtration. The steam condensate generated during the drying process is collected in a centralized drainage tank for testing and discharge. This process does not add any chemical agents or additives (demulsifiers, surfactants, flocculants), only removing suspended solids and excess silica. The reduction process agitates the raw water, accelerating the demulsification of oil in the water and preparing it for the next oil removal process.

[0013] 2. Degreasing

[0014] 2.1 Air flotation for oil removal: A large amount of oil in the wastewater is removed using an air flotation device. The oil then enters a settling tank for further dewatering and recovery. The oil content in the air flotation effluent is ≤50mg / L before proceeding to the next oil removal process.

[0015] 2.2 Dedicated membrane for oil removal

[0016] 2.2.1 Oil-water separation: Ultrafiltration membrane separates oil and water using a two-stage separation process. The final effluent contains ≤0.2 mg / L of oil and ≤0.3 mg / L of TSS. The membrane concentrate is returned to the previous impurity removal process, while the membrane desiccant enters the next process for desiliconization and hardening removal.

[0017] 3. Desilication, removal of heavy metals and sulfate ions

[0018] 3.1 To remove sulfate ions, barium chloride is added to the water after oil removal. The sulfate ions combine with the barium to form barium sulfate precipitate, while the chloride ions remain in the water. The amount of barium chloride added is 0.95-0.98% of the theoretical amount, and the amount of barium chloride is controlled to not exceed the limit.

[0019] 3.2 The pH value is adjusted from 2-3 to 11.1-11.5 using a high-speed mixer. The alkali used for pH adjustment comes from the alkali produced by the electrolysis of sodium chloride in the subsequent process. Heavy metals in the water form hydroxide precipitates, and silicon reacts with excess magnesium in the raw water to form magnesium silicate precipitates.

[0020] 3.3 Solid-liquid separation is achieved by combining a tubular / plate volume reduction filter with a filter press. The purified water from the volume reduction filter enters the next degassing process, while the concentrated water and filter cake washing water from the filter press are returned to the front end of this process for recycling. The filter cake from the filter press is washed, dried, and treated as waste or further recycled. The steam condensate generated during the drying process is collected in a centralized drainage tank for monitoring and discharge.

[0021] 3.4 In this process, barium chloride is added at the theoretical value of 0.95-0.98 during sulfate removal. Barium does not leave residues in the water; it merely replaces sulfate ions with chloride ions, resulting in a decrease in the variety of components in the water, not an increase. Compared to other sulfate removal methods (adding scale inhibitors, defoamers, resin exchange, flocculants), this method is more cost-effective. Subsequent membrane and evaporation crystallization processes do not require the addition of scale inhibitors and defoamers. Under alkaline conditions, silicon combines with magnesium in the raw water to form insoluble magnesium silicate, while other heavy metals form insoluble hydroxides or dehydrate to precipitate oxides, achieving desiliconization and hardening removal. The residual silicon and heavy metals in the water meet the requirements of the subsequent salt separation process.

[0022] 4. Degassing to remove methane, methanol, inorganic ammonia, and small-molecule organic matter from wastewater.

[0023] 4.1 After the wastewater undergoes impurity removal, oil removal, desiliconization and hardening removal, in addition to sodium, potassium, calcium, magnesium and trace amounts of silicon and heavy metals, there are also small molecule organic matter and inorganic ammonia in the water. The organic matter is mainly methane and methanol, which need to be removed before salt separation.

[0024] 4.2 Using a self-developed vacuum microwave degassing tower, alkaline mixed salt water is atomized under negative pressure, and the droplets are heated by microwaves. Methane and inorganic ammonia are desorbed under negative pressure, while methanol in the droplets preferentially desorbs over water via microwave absorption. Other small-molecule organic compounds are also desorbed during the heating process, leaving the water body and entering a gas collection tank to complete degassing. This process is more energy-efficient than traditional steam stripping, as the absence of steam increases the subsequent evaporation rate, and it allows for a large amount of small-molecule organic compounds to be desorbed. The collected mixed gas is then treated in a flare cabinet for harmless disposal.

[0025] 5. Heterogeneous ozone catalytic oxidation

[0026] 5.1 The degassed alkaline mixed brine is further treated with an ozone catalytic oxidation device to remove residual organic matter and ammonia.

[0027] 5.2 The oxidized water is then filtered through an ultrafiltration system. The concentrated ultrafiltration water is returned to the desiliconization and heavy metal removal process, while the ultrafiltration permeate is sent to the next process for salt separation.

[0028] 6. Membrane separation: Membrane separation of monovalent and divalent salts.

[0029] 6.1 Adjust the pH of the mixed brine to 6.5. The acid used to adjust the pH value comes from the hydrochloric acid produced by the electrolysis of sodium chloride in the subsequent process.

[0030] 6.2 Salt separation using a (100 Da) nanofiltration membrane. Control the concentration of monovalent salts to ≥10% and the concentration of divalent salts to ≤30%.

[0031] 7. Evaporation and crystallization

[0032] 7.1 Separation by evaporation and crystallization of monovalent salts.

[0033] 7.1.1 Evaporate monovalent salts to precipitate sodium chloride using an MVR single crystallization evaporation system until potassium chloride, calcium chloride, and magnesium chloride are co-saturated.

[0034] 7.1.2 Sodium chloride is recovered using a horizontal siphon scraper centrifuge. A saturated brine spray pipe is installed inside the centrifuge drum to directly spray the filter cake at full speed, washing away trace amounts of organic mother liquor and ammonium impurities trapped on the crystal surface, thus improving salt quality. The saturated brine is derived from the sodium chloride recovered by evaporation of divalent brine and the monovalent brine separated by the membrane.

[0035] 7.1.3 Dry sodium chloride in a sulfide bed until the moisture content is within acceptable limits.

[0036] 7.1.4 The recovered sodium chloride meets the industry standard T / ZGZS0302-2023.

[0037] 7.1.4 The final mother liquor from the evaporation of monovalent salts enters the divalent salt evaporation and crystallization process for further evaporation and crystallization.

[0038] 7.2 Separation by Evaporation and Crystallization of Divalent Salt Water

[0039] 7.2.1 Separate sodium chloride by evaporating divalent salts using an MVR evaporator until potassium chloride, calcium chloride, and magnesium chloride are co-saturated.

[0040] 7.2.1.1 Sodium chloride is recovered using a horizontal siphon scraper centrifuge. A saturated brine spray pipe is installed inside the centrifuge drum to directly spray the filter cake at full speed, washing away trace amounts of organic mother liquor and ammonium impurities trapped on the crystal surface, thus improving salt quality. The saturated brine is derived from the sodium chloride recovered by evaporation of divalent brine and the monovalent brine separated by the membrane.

[0041] 7.2.1.2 Part of the precipitated sodium chloride is mixed with monovalent brine from the membrane to prepare a saturated sodium chloride cleaning solution, and another part is mixed with monovalent brine from the membrane to prepare an electrolyte. The remaining portion is returned to the monovalent brine from the membrane separation for recrystallization, which only improves the purity of sodium chloride separated by divalent salt, without changing the evaporation rate of monovalent brine.

[0042] 7.2.2 The mother liquor from the sodium chloride separation is then used to crystallize and separate potassium chloride using an MVR evaporator until calcium chloride and magnesium chloride are co-saturated.

[0043] 7.2.2.1 Potassium chloride is recovered using a horizontal centrifuge and dried using a sulfidation bed until the moisture content is within acceptable limits.

[0044] 7.2.2.2 The recovered potassium chloride meets the Class II standard of the National Industrial Salt (GB6549-2011).

[0045] 7.2.3 The final mother liquor is circulated and evaporated using a multi-effect evaporator to crystallize a mixed salt of calcium chloride and magnesium chloride.

[0046] 7.2.3.1 Use a horizontal centrifuge to recover the mixed salts of calcium chloride and magnesium chloride, and dry them with a sulfide bed until the moisture content is qualified.

[0047] 7.2.3.2 The de-icing agent produced meets the national standard GB / T 23851-2017.

[0048] 7.3 Evaporative cooling water,

[0049] 7.3.1 Evaporative cooling water is collected in a centralized drainage pool.

[0050] 7.3.2 The test results meet the GB18918-2002 standard for external discharge.

[0051] 8. Electrolysis of sodium chloride

[0052] 8.1 Prepare an electrolyte solution by evaporating monovalent salt water and sodium chloride obtained by divalent salt crystallization.

[0053] 8.2 Brine purification: chelating resin is used to remove cations to meet the requirements for bipolar membrane electrolysis feed water.

[0054] 8.2.1 After the resin is saturated, it is regenerated with electrolytic hydrochloric acid, and the regenerated water is returned to the divalent salt for evaporation and crystallization.

[0055] 8.3 The qualified brine is subjected to bipolar membrane electrolysis. This electrolysis process directly produces alkali and hydrochloric acid without producing chlorine gas. The produced alkali and hydrochloric acid are used to adjust the pH value in each process, and the electrolysis amount is based on the requirements of each process. After the pH value is adjusted, sodium hydroxide and hydrochloric acid are reduced back to sodium chloride. They only play a condition control role and do not actually participate in the recombination of other elements to generate new substances, thus not changing the total composition of the gas field water treatment process.

[0056] The beneficial effects of this invention are:

[0057] Wastewater from oil and gas fields has been utilized to the greatest extent possible. Only a small amount of solid waste is generated, and four products are obtained: sodium chloride, potassium chloride, calcium chloride / magnesium chloride (snow-resolving agent), and industrial feedstock oil. All water is discharged in compliance with standards.

[0058] The treatment process involves adding only a small amount of barium chloride, eliminating the need for other agents (traditional dual alkalis, demulsifiers, scale inhibitors, defoamers, flocculants, surfactants, etc.), resulting in minimal solid waste and the lowest treatment cost. Energy consumption is solely electricity, and emissions consist only of water. In the Tarim Oilfield of Xinjiang, where green electricity is abundant but water resources are scarce, this advantage is even more pronounced. If the salt recovered through this method is extended into the industrial chain to produce sodium hydroxide and polyvinyl chloride, water treatment costs can be turned from negative to positive, as the underground is not wastewater but a salt lake and a resource. Detailed Implementation

[0059] Example 1

[0060] A 10-ton water sample from the Tarim Oil and Gas Field (A) processing station had a pH of 4.92.

[0061] Experimental preparation: One pilot-scale system for treating gas production wastewater from an oil and gas field.

[0062] Processing capacity: 10t / d

[0063] 1) Connect the wastewater tank to the pilot system feed pipe, and connect the external tap water to the system cleaning water pipe (use evaporation condensate during formal production).

[0064] 2) Auxiliary reagents: 3 kg of sodium hydroxide (caustic soda flakes). Prepare a 10% aqueous solution of sodium hydroxide with pure water and add it to the alkali tank #18 in the system. Electrolytic alkali is used for continuous production and is not purchased externally.

[0065] 3) Prepare 10 kg (30%) of hydrochloric acid into a 15% dilute acid solution and add it to acid tank #17 in the system. Electrolyzed hydrochloric acid will be used for continuous production and will not be purchased externally.

[0066] 4) 50 kg of barium chloride, prepared as a 10% aqueous solution with pure water, is added to the system's No. 19 reagent tank.

[0067] 5) 36 kg of sodium chloride was prepared into a saturated aqueous solution with pure water and added to the backwash tank of a #20 centrifuge. The sodium chloride recovered during production by evaporation using divalent salts is not purchased externally.

[0068] Process:

[0069] 1. Removing impurities

[0070] 1.1 Open the discharge valve of the ton container and start the pilot system. Adjust the feed pump flow rate to 500 L / h, and adjust the No. 1 hydrochloric acid metering pump to ensure the pH value of the water outlet from the No. 1 high-speed mixer is 2-3, precisely adjusted to 2.6. Unqualified water from the front end is returned to the ton container. Qualified water enters the No. 1 buffer tank of the system, with a maximum height of 1.5 m³. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0071] 1.2 The system's No. 1 buffer tank discharge pump automatically starts at the upper limit time, adjusting the flow rate to 500 L / h at the No. 1 volumetric filter outlet. The clean water from the volumetric filter enters the flotation tank, while the concentrated water enters the No. 1 filter press. The effluent from the No. 1 filter press returns to the No. 1 buffer tank for circulation. The filter press has a filtration area of ​​20 m². 2 .

[0072] 1.3 Impurity removal complete.

[0073] 1.3.1 Hydrochloric acid dosage: 7.32 L; Filter press wet cake: 15.6 kg, dried cake: 3.52 kg

[0074] Use 20L of tap water to clean the filter cake, then return it to the container and enter the system.

[0075] 1.3.2 System increment: 7.32L hydrochloric acid, 20kg tap water, 15.6kg filter cake from filter press, the actual system increment is 11.72kg.

[0076] 1.3.3 TSS of impurity-removed effluent: 8.2 mg / L

[0077] 2. Degreasing

[0078] 2.1 The volume of the dissolved air flotation tank is 1.0 m³. 3 ,

[0079] When the deionized water reaches 80% of the dissolved air flotation (DAF) tank's volume, the air compressor starts, the dissolved air tank's circulation pump starts, and the DAF tank begins operation. When the DAF tank reaches its upper limit and the clear water chamber is full, the discharge pump automatically starts, and clear water enters the No. 2 buffer tank, which has a maximum capacity of 0.5m. 3 Lower limit 0.1 m 3 The floating oil in the flotation tank is scraped into the oil collection chamber and then pumped by a gear pump to the settling tank for further oil separation. When the floating oil in the settling tank reaches the upper limit, the oil is automatically discharged to the recovery tank for recycling. When the settling water reaches the upper limit, it is pumped back to the flotation tank for circulation.

[0080] 2.1.1 The oil content in the air flotation effluent was 8.6 mg / L.

[0081] 2.1.2 3.57 kg of oil was recovered after the air flotation process.

[0082] 2.2 Dedicated membrane for oil removal

[0083] 2.2.1 The oil removal membrane in the system is WKXZ-004 / WKXL-200. When the No. 2 buffer tank reaches its upper limit, the membrane feed pump automatically starts, allowing water to enter the oil removal membrane. The permeate flow rate of the oil removal membrane is adjusted to 500 L / h. The concentrate from the oil removal membrane is returned to the No. 1 buffer tank for circulation, and the membrane permeate enters the No. 3 buffer tank. The upper limit of the buffer tank is 1.0 m³. 3 Lower limit 0.5 m 3 .

[0084] 2.2.2 The oil film removal product water contained 0.18 mg / L of oil and 0.83 mg / L of TSS.

[0085] 2.2.3 Table of water composition after oil removal membrane

[0086] 3. Desiliconization to remove heavy metals and sulfate ions

[0087] 3.1 When the liquid level in system #3 buffer tank reaches the upper limit, the discharge pump of #3 buffer tank automatically starts, adjusting the discharge flow rate to 500 L / h at the outlet flow rate of #2 high-speed mixer. After the mixer outlet flow rate stabilizes, the barium chloride metering pump is started, and the flow rate is adjusted to 7.65 L / h. The degreased water with added barium chloride enters #4 buffer tank, with a maximum upper limit of 1.5 m. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0088] 3.2 When the upper limit of the system's No. 4 buffer tank is reached, the discharge pump automatically starts. Adjust the No. 4 alkali metering pump to ensure the pH value of the water outlet from the No. 3 high-speed mixer is between 11.1 and 11.5, precisely adjusted to 11.3. Unqualified water from the front end returns to the No. 4 buffer tank. Qualified water enters the system's No. 5 buffer tank, with a maximum capacity of 1.5m. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0089] 3.3 When the liquid level in system #5 buffer tank reaches the upper limit, the discharge pump automatically starts, adjusting the flow rate to 500 L / h at the outlet of #2 reduction filter. Clean water from #2 reduction filter enters #6 buffer tank, with the upper limit of the buffer tank at 0.5 m. 3 Lower limit 0.1 m 3 The concentrated wastewater from filter press #2 enters the filter press. The effluent from filter press #2 is returned to the buffer tank #4 for recirculation. The filter press has a filtration area of ​​40 m². 2 .

[0090] 3.4 Silicon and heavy metal removal process completed

[0091] 3.4.1 Alkali dosage: 18.5L; Filter press wet cake yield: 118.3kg; Drying yield: 28.4kg;

[0092] 200 kg of tap water is used to clean the filter cake, and the cleaned water is circulated into buffer tank #4.

[0093] 1.3.2 System increments: alkali 18.5L, tap water 200kg, filter cake from filter press 118.3kg, barium chloride 153L.

[0094] 1.3.3 Water composition table after silicon and heavy metal removal

[0095] 4. Degassing

[0096] 4.1 When the water level in buffer tank #6 reaches the lower limit, the vacuum pump of the degassing tower starts, and the microwave power supply's automatic program begins operation. When the water level in buffer tank #6 reaches the upper limit, the discharge pump and pipeline booster pump start, and the degassing tower begins full operation. The degassing tower inlet flow rate is adjusted to 500 L / h, the microwave power supply is 10 kW, and the actual operating power is 7.35 kW. The removed ammonia, methane, methanol, and organic matter are dehydrated by molecular sieves and then enter the flare cabinet for combustion. The degassed water is pumped to buffer tank #7, with a maximum upper limit of 0.5 m. 3 Lower limit 0.1m 3 Stir at 130 rpm.

[0097] 5. Heterogeneous ozone catalytic oxidation

[0098] 5.1 When the influent level of the No. 7 buffer tank in system 5.1 reaches the upper limit, the discharge pump starts, and the flow rate is adjusted to 500L / h of effluent from the oxidation system. The oxidation system model is FW14-001, with an installed power of 3.5kW, an operating power of 3.3kW, and a processing capacity of 1.0m³. 3 / h. The oxidized water enters buffer tank #8, which has a maximum capacity of 0.5m. 3 Lower limit 0.1m 3 .

[0099] 5.2 When the influent level in buffer tank #8 reaches its upper limit, the discharge pump starts, and the flow rate is adjusted to 500 L / h of ultrafiltration membrane permeate. The ultrafiltration membrane material is PTFE. The ultrafiltration membrane concentrate returns to buffer tank #5, and the ultrafiltration membrane permeate enters buffer tank #9, with a top limit of 0.5 m. 3 Lower limit 0.1m 3 .

[0100] 5.3 Ultrafiltration Permeate Composition Table

[0101] 6. Membrane separation to separate monovalent and divalent salts.

[0102] 6.1 When the influent level in system #9 buffer tank reaches the upper limit, the discharge pump starts, and the flow rate is adjusted to 500 L / h at the outlet of system #4 high-speed mixer. The #9 hydrochloric acid metering pump is adjusted to ensure the pH of the water outlet from system #4 is 6-7, precisely adjusted to 6.5. Unqualified water is returned to system #9 buffer tank. Qualified water enters system #10 buffer tank, with a maximum capacity of 1.5m. 3 Lower limit 0.5 m, residence time 2 hours, stirring 15 r / min.

[0103] 6.1.1 System increment: Hydrochloric acid usage 5.36L.

[0104] 6.2% salt content

[0105] 6.2.1 Membrane type: 100Da nanofiltration membrane and RO membrane

[0106] 6.2.2 Salt Classification: Monovalent salt concentration ≥ 10%, divalent salt concentration ≤ 30%

[0107] 6.2.3 When the influent level of system 10# buffer tank reaches the upper limit, the discharge will begin. Adjust the flow rate to 500 L / h for the membrane desalination system influent. Monovalent brine enters 11# buffer tank, and divalent brine enters 12# buffer tank. The upper limit of the buffer tanks is 2.5 m. 3 Lower limit 0.5m 3 .

[0108] 6.2.4 Salt composition design table

[0109] 6.2.5.1 Monovalent Salts

[0110] 6.2.5.2 Divalent Salts

[0111] 6.2.6 Monovalent salt concentration 12.42%, divalent salt concentration 19.28%, no RO concentrate.

[0112] 7. Evaporation and crystallization

[0113] 7.1 Evaporation and crystallization of monovalent salts

[0114] 7.1.1 When the influent level in the No. 11 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, feeding the No. 1 MVR evaporator. The feed flow rate is adjusted to 376.65 L / h, and the evaporation rate is controlled at 200 kg / h. The evaporation rate of the No. 2 MVR evaporator is controlled at 129.32 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 1 crystallizer. The upper limit of the No. 1 crystallizer is 2.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0115] 7.1.2 When the upper limit of crystallizer #1 is reached, the discharge valve of crystallizer #1 is opened, and centrifuge #1 (GKH630-N) starts working, with the mother liquor entering buffer tank #12. When the centrifuge is full, the centrifuge backwash pump starts, rinsing the sodium chloride with saturated brine, and the rinse water enters buffer tank #12. After rinsing is complete, the centrifuge discharges the sodium chloride to the system's #1 sulfurized bed for drying and packaging, completing the sodium chloride recovery.

[0116] 7.1.3 Evaporation and crystallization of monovalent salts are completed.

[0117] 7.1.3.1 System 15 # Storage Tank Condensate: 6257 kg

[0118] 888 kg of sodium chloride was recovered.

[0119] Use 85L of saturated brine for cleaning.

[0120] 7.1.3.2 Total Composition Table of Mother Liquor from Sodium Chloride Crystallization Centrifugation

[0121] 7.2 Evaporation and crystallization of divalent salts

[0122] 7.2.1 Design table of composition of divalent concentrated water before evaporation

[0123] 7.2.2 Separation of sodium chloride

[0124] 7.2.2.1 When the influent level in the system's No. 12 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, and feed is introduced into the No. 3 MVR evaporator. The feed flow rate is adjusted to 111.71 L / h, and the evaporation rate is controlled at 69.26 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 2 crystallizer. The upper limit of the No. 2 crystallizer is 2.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0125] 7.2.2.2 When the upper limit of crystallizer #2 is reached, the discharge valve of crystallizer #2 is opened, and centrifuge #2 (GK450-N) is started. When the centrifuge is full, the centrifuge backwash pump is started, and saturated brine is used to wash the sodium chloride. The washing water enters buffer tank #12. The centrifuge discharges the material into the sodium chloride storage silo (in industrial production, this sodium chloride is used to prepare saturated washing water and electrolyte; the excess is dissolved in monovalent brine for circulating evaporation and crystallization). The centrifugal mother liquor enters buffer tank #13. The upper limit of buffer tank #13 is 2.5m. 3 Lower limit 0.5 m 3 .

[0126] 7.2.2.3 Separation complete.

[0127] 7.2.2.3.1 System 15 # Storage Tank Condensate: 1316 kg

[0128] 212.76 kg of sodium chloride was recovered, and 23 L of saturated brine was used for washing.

[0129] 7.2.2.3.2 Total Composition Table of Mother Liquor from Sodium Chloride Crystallization Centrifugation

[0130] 7.2.3 Separation of potassium chloride

[0131] 7.2.3.1 When the influent level in the system's No. 13 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, and feed is introduced into the No. 4 MVR evaporator. The feed flow rate is adjusted to 38.43 L / h, and the evaporation rate is controlled at 13.26 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 3 crystallizer. The upper limit of the No. 3 crystallizer is 1.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0132] 7.2.3.2 With the upper limit of crystallizer #3 set, the discharge valve of crystallizer #3 is opened, and centrifuge #3 (GK400-N) is started. The centrifuge discharges material into the potassium chloride #2 sulfurized bed for drying and packaging. The centrifuged mother liquor enters the buffer tank #14. The upper limit of buffer tank #14 is 0.3m. 3 Lower limit 0.1m 3 .

[0133] 7.2.3.3 Separation complete.

[0134] 7.2.3.3.1 System 15 # Storage Tank Condensate 252kg,

[0135] 88.06 kg of potassium chloride was recovered.

[0136] 7.2.3.3.2 Total Composition Table of Potassium Chloride Crystallization Centrifugation Mother Liquor

[0137] 7.2.4 Recovery of calcium chloride and magnesium chloride

[0138] 7.2.4.1 When the influent level in the system's No. 14 buffer tank reaches the lower limit, the evaporator system starts, and the multi-effect evaporator preheats. When the influent level reaches the upper limit, the forced circulation pump starts, and the multi-effect evaporator is fed and circulated. The circulation flow rate is adjusted to 25.8 L / h, and the evaporation rate is controlled at 13.6 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 4 crystallizer. The upper limit of the crystallizer is 0.3 m. 3 Lower limit 0.1m 3 Stir at 3 R / min.

[0139] 7.2.4.2 The upper limit of crystallizer #4 in the system is reached, the discharge valve of crystallizer #4 is opened, and centrifuge #4 (GK400-N) is started. The centrifuge discharges material to the calcium chloride and magnesium chloride #3 sulfurized bed for drying and packaging. The mother liquor from the centrifugation enters the buffer tank #14 for circulation evaporation.

[0140] 7.2.4.3 Separation complete.

[0141] 7.2.4.3.1 System 15 # Storage Tank Condensate 285.2kg,

[0142] 334.15 kg of mixed salt of calcium chloride and magnesium chloride were recovered.

[0143] 8. After processing 10 tons of water samples, a total of [amount] was recovered.

[0144] 8.1 Sodium chloride 1100.76 kg,

[0145] 8.2 kJ of potassium chloride,

[0146] 8.3 Calcium chloride, magnesium chloride 334.15

[0147] 8.4 Evaporation condensate 8110.2kg

[0148] 8.5 kg of crude oil was recovered.

[0149] 8.6 Use 18.5L of 10% alkali and 153L of 10% barium chloride solution.

[0150] 8.7 Use 12.68L of 15% acid and 108L of saturated saline solution.

[0151] 8.8 Solid waste 31.92 kg

[0152] 8.9 Quality Table of Recycled Products

[0153] 8.9.1 Sodium chloride

[0154] 8.9.2 Potassium chloride

[0155] 8.9.3 Mixed salts of calcium chloride and magnesium chloride (snow-dissolving agent)

[0156] 8.9.4 Evaporation condensate

[0157] 9. Electrolysis of sodium chloride and preparation of saturated brine

[0158] 9.1 Electrolyzed brine

[0159] 9.1.1 Dissolve 6.9 kg of sodium chloride evaporated from divalent brine into a 60 L membrane-separated monovalent brine solution.

[0160] 9.1.2 Add the prepared brine to the system brine storage tank and turn on the discharge pump to control the flow rate at 2.5 BV. The brine passes through the resin column and enters the #16 electrolytic buffer tank.

[0161] 9.1.3 Start the bipolar membrane electrolysis system. The acid and alkali produced by the bipolar membrane electrolysis of hard water in the #16 electrolysis buffer tank are respectively sent to the #17 acid tank and the #18 alkali tank.

[0162] 9.1.4 The process involves separate reflux electrolysis of acid and alkali, controlling the sodium hydroxide concentration in tank #18 to 10% and the hydrochloric acid concentration in tank #17 to 15% before completion.

[0163] 9.1.5 The volume of acid tank #17 in the system increases by 29.23L, and the volume of alkali tank #18 increases by 30.77L.

[0164] 9.2 Preparation of saturated saline solution

[0165] 9.2.1 Dissolve 23.43 kg of sodium chloride evaporated from divalent brine into 109.36 L of monovalent brine separated by membrane separation.

[0166] 9.2.2 Add the prepared saturated brine to the backwash tank of centrifuge #20.

[0167] Example 2

[0168] A 10-ton water sample from the Tarim Oil and Gas Field (B) processing station had a pH of 6.07.

[0169] Process:

[0170] 1. Removing impurities

[0171] 1.1 Open the discharge valve of the ton container and start the pilot system. Adjust the feed pump flow rate to 500 L / h, and adjust the No. 1 hydrochloric acid metering pump to ensure the pH value of the water outlet from the No. 1 high-speed mixer is 2-3, precisely adjusted to 2.6. Unqualified water from the front end is returned to the ton container. Qualified water enters the No. 1 buffer tank of the system, with a maximum height of 1.5 m³. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0172] 1.2 The system's No. 1 buffer tank discharge pump automatically starts at the upper limit time, adjusting the flow rate to 500 L / h at the No. 1 reduction filter's clean water outlet. Clean water from the tubular filter enters the flotation tank, while concentrated water enters the No. 1 filter press. The filter press effluent returns to the No. 1 buffer tank for circulation. The No. 1 filter press has a filtration area of ​​20 m². 2 .

[0173] 1.3 Impurity removal complete.

[0174] 1.3.1 Hydrochloric acid dosage: 5.72 L / kg; 8.9 kg of wet filter cake from the filter press, 2.6 kg after drying.

[0175] 20kg of tap water is used to clean the filter cake, and the water is returned to the ton container and then into the system.

[0176] 1.3.3 TSS of impurity-removed effluent: 7.3 mg / L

[0177] 2. Degreasing

[0178] 2.1 The volume of the dissolved air flotation tank is 1.0 m³. 3 ,

[0179] When the deionized water reaches 80% of the dissolved air flotation (DAF) tank's volume, the air compressor starts, the dissolved air tank's circulation pump starts, and the DAF tank begins operation. When the DAF tank reaches its upper limit and the clear water chamber is full, the discharge pump automatically starts, and clear water enters the No. 2 buffer tank, which has a maximum capacity of 0.5m. 3 Lower limit 0.1 m 3 The floating oil in the flotation tank is scraped into the oil collection chamber and then pumped by a gear pump to the settling tank for further oil separation. When the floating oil in the settling tank reaches the upper limit, the oil is automatically discharged to the recovery tank for recycling. When the settling water reaches the upper limit, it is pumped back to the flotation tank for circulation.

[0180] 2.1.1 The oil content in the air flotation effluent was 6.3 mg / L.

[0181] 2.1.2 2.22 kg of oil was recovered after the air flotation process ended.

[0182] 2.2 Dedicated membrane for oil removal

[0183] 2.2.1 The oil removal membrane in the system is WKXZ-004 / WKXL-200. When the No. 2 buffer tank reaches its upper limit, the membrane feed pump automatically starts, allowing water to enter the oil removal membrane. The permeate flow rate of the oil removal membrane is adjusted to 500 L / h. The concentrate from the oil removal membrane is returned to the No. 1 buffer tank for circulation, and the membrane permeate enters the No. 3 buffer tank. The upper limit of the buffer tank is 1.0 m³. 3 Lower limit 0.5 m 3 .

[0184] 2.2.2 The oil film removal product water contains 0.15 mg / L of oil and 0.72 mg / L of TSS.

[0185] 2.2.3 Table of water composition after oil removal membrane

[0186] 3. Desiliconization to remove heavy metals and sulfate ions

[0187] 3.1 When the liquid level in system #3 buffer tank reaches the upper limit, the discharge pump of #3 buffer tank automatically starts, adjusting the discharge flow rate to 500 L / h at the outlet flow rate of #2 high-speed mixer. After the mixer outlet flow rate stabilizes, the barium chloride metering pump is started, and the flow rate is adjusted to 12.1 L / h. The degreased water with added barium chloride enters #4 buffer tank, with a maximum upper limit of 1.5 m. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0188] 3.2 When the upper limit of the system's No. 4 buffer tank is reached, the discharge pump automatically starts. Adjust the No. 4 alkali metering pump to ensure the pH value of the water outlet from the No. 3 high-speed mixer is between 11.1 and 11.5, precisely adjusted to 11.3. Unqualified water from the front end returns to the No. 4 buffer tank. Qualified water enters the system's No. 5 buffer tank, with a maximum capacity of 1.5m. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0189] 3.3 When the liquid level in system #5 buffer tank reaches the upper limit, the discharge pump automatically starts, adjusting the flow rate to 500L / h at the outlet of #2 reduction filter. Clean water from the tubular filter enters #6 buffer tank, with the upper limit of the buffer tank being 0.5m. 3 Lower limit 0.1 m 3 The concentrated wastewater enters filter press #2. The effluent from filter press #2 is returned to buffer tank #4 for recirculation. The filter press has a filtration area of ​​40 m². 2 .

[0190] 3.4 Silicon and heavy metal removal process completed

[0191] 3.4.1 Alkali dosage: 16.91L; Filter press wet cake yield: 138.45kg, drying yield: 27.7kg; Barium chloride dosage: 242.3L

[0192] In Example 1, 200 kg of evaporative cooling water was used to clean the filter cake. The cleaned water was then circulated into buffer tank #4.

[0193] 1.3.3 Water composition table after silicon and heavy metal removal

[0194] 4. Degassing

[0195] 4.1 When the water level in buffer tank #6 reaches the lower limit, the vacuum pump of the degassing tower starts, and the microwave power supply's automatic program begins operation. When the water level in buffer tank #6 reaches the upper limit, the discharge pump and pipeline booster pump start, and the degassing tower begins full operation. The degassing tower inlet flow rate is adjusted to 500 L / h, the microwave power supply is 10 kW, and the actual operating power is 7.35 kW. The removed ammonia, methane, methanol, and organic matter are dehydrated by molecular sieves and then enter the flare cabinet for combustion. The degassed water is pumped to buffer tank #7, with a maximum upper limit of 0.5 m. 3Lower limit 0.1m 3 Stir at 130 rpm.

[0196] 5. Heterogeneous ozone catalytic oxidation

[0197] 5.1 When the influent level of the No. 7 buffer tank in system 5.1 reaches the upper limit, the discharge pump starts, and the flow rate is adjusted to 500L / h of effluent from the oxidation system. The oxidation system model is FW14-001, with an installed power of 3.5kW, an operating power of 3.3kW, and a processing capacity of 1.0m³. 3 / h. The oxidized water enters buffer tank #8, which has a maximum capacity of 0.5m. 3 Lower limit 0.1m 3 .

[0198] 5.2 When the influent level in buffer tank #8 reaches its upper limit, the discharge pump starts, and the flow rate is adjusted to 500 L / h of ultrafiltration membrane permeate. The ultrafiltration membrane material is PTFE. The ultrafiltration membrane concentrate returns to buffer tank #5, and the ultrafiltration membrane permeate enters buffer tank #9, with a top limit of 0.5 m. 3 Lower limit 0.1m 3 .

[0199] 5.3 Ultrafiltration Permeate Composition Table

[0200] 6. Membrane separation to separate monovalent and divalent salts.

[0201] 6.1 When the influent level in system #9 buffer tank reaches the upper limit, the discharge pump starts, and the flow rate is adjusted to 500 L / h at the outlet of system #4 high-speed mixer. The #9 hydrochloric acid metering pump is adjusted to ensure the pH of the water outlet from system #4 is 6-7, precisely adjusted to 6.5. Unqualified water is returned to system #9 buffer tank. Qualified water enters system #10 buffer tank, with a maximum capacity of 1.5m. 3 Lower limit 0.5 m, residence time 2 hours, stirring 15 r / min.

[0202] 6.1.1 System increment: Hydrochloric acid dosage 4.55L.

[0203] 6.2% salt content

[0204] 6.2.1 Membrane type: 100Da nanofiltration membrane and RO membrane

[0205] 6.2.2 Salt Classification: Monovalent salt concentration ≥ 10%, divalent salt concentration ≤ 30%

[0206] 6.2.1 Membrane type: 100Da nanofiltration membrane and RO membrane

[0207] 6.2.3 When the influent level of the system's No. 10 buffer tank reaches its upper limit, the discharge will begin. Adjust the flow rate to 500 L / h for the membrane desalination system influent. Adjust the membrane desalination system parameters until the circulating desalination reaches 80% monovalent salt and 20% divalent salt. Monovalent brine enters the No. 11 buffer tank, and divalent brine enters the No. 12 buffer tank. The upper limit of the buffer tanks is 2.5 m. 3 Lower limit 0.5 m 3 .

[0208] 6.2.4 Salt composition design table

[0209] 6.2.5.1 Monovalent Salts

[0210] 6.2.5.2 Divalent Salts

[0211] 6.2.6 The concentration of monovalent salts was 11.53%, and the concentration of divalent salts was 21.23%.

[0212] 6.2.7 RO Concentrated Water 3280.55L

[0213] 7. Evaporation and crystallization

[0214] 7.1 Evaporation and crystallization of monovalent salts

[0215] 7.1.1 When the influent level in the system's No. 11 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, feeding the No. 1 MVR evaporator. The feed flow rate is adjusted to 258.87 L / h, controlling the evaporation rate at 203.73 kg / h, and the evaporation rate of the No. 2 MVR evaporator is controlled at 34.37 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 1 crystallizer. The upper limit of the No. 1 crystallizer is 2.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0216] 7.1.2 When the upper limit of crystallizer #1 is reached, the discharge valve of crystallizer #1 is opened, and centrifuge #1 (GKH630-N) starts working, with the mother liquor entering buffer tank #12. When the centrifuge is full, the centrifuge backwash pump starts, rinsing the sodium chloride with saturated brine, and the rinse water enters buffer tank #12. After rinsing is complete, the centrifuge discharges the sodium chloride to the system's #1 sulfurized bed for drying and packaging, completing the sodium chloride recovery.

[0217] 7.1.3 Evaporation and crystallization of monovalent salts are completed.

[0218] 7.1.3.1 System 15 # Storage Tank Condensate: 4523.73 kg

[0219] 598.7 kg of sodium chloride was recovered.

[0220] Use 75 kg of saturated brine for cleaning.

[0221] 7.1.3.2 Total Composition Table of Mother Liquor from Sodium Chloride Crystallization Centrifugation

[0222] 7.2 Evaporation and crystallization of divalent salts

[0223] 7.2.1 Design table of composition of divalent concentrated water before evaporation

[0224] 7.2.2 Separation of sodium chloride

[0225] 7.2.2.1 When the influent level in the system's No. 12 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, and feed is introduced into the No. 3 MVR evaporator. The feed flow rate is adjusted to 57.15 L / h, and the evaporation rate is controlled at 36.7 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 2 crystallizer. The upper limit of the No. 2 crystallizer is 2.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0226] 7.2.2.2 With the upper limit of crystallizer #2 set, the discharge valve of crystallizer #2 is opened, and centrifuge #2 (GK450-N) is started. The centrifuge discharges material into the sodium chloride storage silo (in industrial production, this sodium chloride is used to prepare saturated cleaning water and electrolyte; the excess is dissolved in monovalent brine for circulating evaporation and crystallization). The centrifuged mother liquor enters buffer tank #13. The upper limit of buffer tank #13 is 2.5m. 3 Lower limit 0.5m 3 .

[0227] 7.2.2.3 Separation complete.

[0228] 7.2.2.3.1 System 15 # Storage Tank Condensate: 697.21 kg

[0229] 154.15 kg of sodium chloride was recovered.

[0230] 7.2.2.3.2 Total Composition Table of Mother Liquor from Sodium Chloride Crystallization Centrifugation

[0231] 7.2.3 Separation of potassium chloride

[0232] 7.2.3.1 When the influent level in the system's No. 13 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, and feed is introduced into the No. 4 MVR evaporator. The feed flow rate is adjusted to 20.12 L / h, and the evaporation rate is controlled at 7.08 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 3 crystallizer. The upper limit of the No. 3 crystallizer is 1.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0233] 7.2.3.2 With the upper limit of crystallizer #3 set, the discharge valve of crystallizer #3 is opened, and centrifuge #3 (GK400-N) is started. The centrifuge discharges material into the potassium chloride #2 sulfurized bed for drying and packaging. The centrifuged mother liquor enters the buffer tank #14. The upper limit of buffer tank #14 is 0.3m. 3 Lower limit 0.1m 3 .

[0234] 7.2.3.3 Separation complete. 7.2.3.3.1

[0236] System 15 storage tank discharged 134.53 kg of condensate and recovered 46.64 kg of potassium chloride.

[0237] 7.2.3.3.2 Total Composition Table of Potassium Chloride Crystallization Centrifugation Mother Liquor

[0238] 7.2.4 Recovery of calcium chloride and magnesium chloride

[0239] 7.2.4.1 When the influent level in the system's No. 14 buffer tank reaches the lower limit, the evaporator system starts, and the multi-effect evaporator preheats. When the influent level reaches the upper limit, the forced circulation pump starts, and the multi-effect evaporator is fed and circulated. The circulation flow rate is adjusted to 11.55 L / h, and the evaporation rate is controlled at 7.22 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 4 crystallizer. The upper limit of the crystallizer is 0.3 m. 3 Lower limit 0.1m 3 Stir at 3 R / min.

[0240] 7.2.4.2 System #4 crystallizer upper limit, #4 crystallizer discharge valve open, #4 centrifuge (GK400-N) start, centrifuge discharges material into calcium chloride and magnesium chloride #3 sulfidation bed.

[0241] Drying and packaging.

[0242] 7.2.4.3 Separation complete.

[0243] 7.2.4.3.1 System 15 # Storage Tank Condensate: 137.23 kg

[0244] 169.23 kg of a mixed salt of calcium chloride and magnesium chloride was recovered.

[0245] 8. After processing 10 tons of water samples, a total of [amount] was recovered.

[0246] 8.1 Sodium chloride 752.85 kg,

[0247] 8.2 Potassium chloride 46.64 kg,

[0248] 8.3 Calcium chloride and magnesium chloride 169.23 kg

[0249] 8.4 Evaporation condensate: 8772.69 kg

[0250] 8.5% recovery rate: 2.22 kg crude oil

[0251] 8.6 Use 16.91 L of 10% alkali and 242.3 L of 10% barium chloride solution.

[0252] 8.7 Use 10.27L of 15% acid and 75kg of saturated brine.

[0253] 8.8 Solid waste 30.3 kg

[0254] 8.9 Quality Table of Recycled Products

[0255] 8.9.1 Sodium chloride

[0256] 8.9.2 Potassium chloride

[0257] 8.9.3 Mixed salts of calcium chloride and magnesium chloride (snow-dissolving agent)

[0258] 8.9.4 Evaporation condensate

[0259] Example 3

[0260] A 10-ton water sample from the Tarim Oil and Gas Field (C) processing station had a pH of 7.7.

[0261] Process:

[0262] 1. Removing impurities

[0263] 1.1 Open the discharge valve of the ton container and start the pilot system. Adjust the feed pump flow rate to 500 L / h, and adjust the No. 1 hydrochloric acid metering pump to ensure the pH value of the water outlet from the No. 1 high-speed mixer is 2-3, precisely adjusted to 2.6. Unqualified water from the front end is returned to the ton container. Qualified water enters the No. 1 buffer tank of the system, with a maximum height of 1.5 m³. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0264] 1.2 The system's No. 1 buffer tank discharge pump automatically starts at the upper limit time, adjusting the flow rate to 500 L / h at the No. 1 reduction filter's clean water outlet. Clean water from the tubular filter enters the flotation tank, while concentrated water enters the No. 1 filter press. The effluent from the No. 1 filter press returns to the No. 1 buffer tank for circulation. The filter press has a filtration area of ​​20 m². 2 .

[0265] 1.3 Impurity removal complete.

[0266] 1.3.1 Hydrochloric acid dosage: 6.1 L / kg; 10 kg of wet filter cake from the filter press, 2.1 kg after drying.

[0267] 10kg of tap water is used to clean the filter cake, and the water is returned to the container and then into the system.

[0268] 1.3.3 TSS of impurity-removed effluent: 8.2 mg / L

[0269] 2. Degreasing

[0270] 2.1 The volume of the dissolved air flotation tank is 1.0 m³. 3 ,

[0271] When the deionized water reaches 80% of the dissolved air flotation (DAF) tank's volume, the air compressor starts, the dissolved air tank's circulation pump starts, and the DAF tank begins operation. When the DAF tank reaches its upper limit and the clear water chamber is full, the discharge pump automatically starts, and clear water enters the No. 2 buffer tank, which has a maximum capacity of 0.5m. 3 Lower limit 0.1 m 3 The floating oil in the flotation tank is scraped into the oil collection chamber and then pumped by a gear pump to the settling tank for further oil separation. When the floating oil in the settling tank reaches the upper limit, the oil is automatically discharged to the recovery tank for recycling. When the settling water reaches the upper limit, it is pumped back to the flotation tank for circulation.

[0272] 2.1.1 The oil content in the effluent from air flotation was 8.4 mg / L.

[0273] 2.1.2 3.53 kg of oil was recovered after the air flotation process ended.

[0274] 2.2 Dedicated membrane for oil removal

[0275] 2.2.1 The oil removal membrane in the system is WKXZ-004 / WKXL-200. When the No. 2 buffer tank reaches its upper limit, the membrane feed pump automatically starts, allowing water to enter the oil removal membrane. The permeate flow rate of the oil removal membrane is adjusted to 500 L / h. The concentrate from the oil removal membrane is returned to the No. 1 buffer tank for circulation, and the membrane permeate enters the No. 3 buffer tank. The upper limit of the buffer tank is 1.0 m³. 3 Lower limit 0.5 m 3 .

[0276] 2.2.2 The oil film removal product water contains 0.15 mg / L of oil and 0.72 mg / L of TSS.

[0277] 2.2.3 Table of water composition after oil removal membrane

[0278] 3. Desiliconization to remove heavy metals and sulfate ions

[0279] 3.1 When the liquid level in system #3 buffer tank reaches the upper limit, the discharge pump of #3 buffer tank automatically starts, adjusting the discharge flow rate to 500 L / h at the outlet flow rate of #2 high-speed mixer. After the mixer outlet flow rate stabilizes, the barium chloride metering pump is started, and the flow rate is adjusted to 3.56 L / h. The degreased water with added barium chloride enters #4 buffer tank, with a maximum upper limit of 1.5 m. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0280] 3.2 When the upper limit of the system's No. 4 buffer tank is reached, the discharge pump automatically starts. Adjust the No. 4 alkali metering pump to ensure the pH value of the water outlet from the No. 3 high-speed mixer is between 11.1 and 11.5, precisely adjusted to 11.3. Unqualified water from the front end returns to the No. 4 buffer tank. Qualified water enters the system's No. 5 buffer tank, with a maximum capacity of 1.5m. 3 Lower limit 0.5 m 3 Residence time 2 hours, stirring 15 r / min.

[0281] 3.3 When the liquid level in system #5 buffer tank reaches the upper limit, the discharge pump automatically starts, adjusting the flow rate to 500L / h at the outlet of #2 reduction filter. Clean water from the tubular filter enters #6 buffer tank, with the upper limit of the buffer tank being 0.5m. 3 Lower limit 0.1 m 3 The concentrated wastewater enters filter press #2. The effluent from the filter press is returned to the buffer tank #4 for recirculation. The filter press has a filtration area of ​​40 m². 2 .

[0282] 3.4 Silicon and heavy metal removal process completed

[0283] 3.4.1 Alkali dosage: 12.3L; Filter press wet cake yield: 71.3kg, drying yield: 8.15kg; Barium chloride dosage: 84.2L

[0284] In Example 2, 100 kg of evaporative condensate was used to clean the filter cake. The cleaned water was then circulated into buffer tank #4.

[0285] 1.3.3 Water composition table after silicon and heavy metal removal

[0286] 4. Degassing

[0287] 4.1 When the water level in buffer tank #6 reaches the lower limit, the vacuum pump of the degassing tower starts, and the microwave power supply's automatic program begins operation. When the water level in buffer tank #6 reaches the upper limit, the discharge pump and pipeline booster pump start, and the degassing tower begins full operation. The degassing tower inlet flow rate is adjusted to 500 L / h, the microwave power supply is 10 kW, and the actual operating power is 7.35 kW. The removed ammonia, methane, methanol, and organic matter are dehydrated by molecular sieves and then enter the flare cabinet for combustion. The degassed water is pumped to buffer tank #7, with a maximum upper limit of 0.5 m. 3 Lower limit 0.1m 3 Stir at 130 rpm.

[0288] 5. Heterogeneous ozone catalytic oxidation

[0289] 5.1 When the influent level of the No. 7 buffer tank in system 5.1 reaches the upper limit, the discharge pump starts, and the flow rate is adjusted to 500L / h of effluent from the oxidation system. The oxidation system model is FW14-001, with an installed power of 3.5kW, an operating power of 3.3kW, and a processing capacity of 1.0m³. 3 / h. The oxidized water enters buffer tank #8, which has a maximum capacity of 0.5m. 3 Lower limit 0.1m 3 .

[0290] 5.2 When the influent level in buffer tank #8 reaches its upper limit, the discharge pump starts, and the flow rate is adjusted to 500 L / h of ultrafiltration membrane permeate. The ultrafiltration membrane material is PTFE. The ultrafiltration membrane concentrate returns to buffer tank #5, and the ultrafiltration membrane permeate enters buffer tank #9, with a top limit of 0.5 m. 3 Lower limit 0.1m 3 .

[0291] 5.3 Ultrafiltration Permeate Composition Table

[0292] 6. Membrane separation to separate monovalent and divalent salts.

[0293] 6.1 When the influent level in system #9 buffer tank reaches the upper limit, the discharge pump starts, and the flow rate is adjusted to 500 L / h at the outlet of system #4 high-speed mixer. The #9 hydrochloric acid metering pump is adjusted to ensure the pH of the water outlet from system #4 is 6-7, precisely adjusted to 6.5. Unqualified water is returned to system #9 buffer tank. Qualified water enters system #10 buffer tank, with a maximum capacity of 1.5m. 3Lower limit 0.5 m, residence time 2 hours, stirring 15 r / min.

[0294] 6.1.1 Hydrochloric acid dosage: 4.62 L.

[0295] 6.2% salt content

[0296] 6.2.1 Membrane type: 100Da nanofiltration membrane and RO membrane

[0297] 6.2.2 Salt Classification: Monovalent salt concentration ≥ 10%, divalent salt concentration ≤ 30%

[0298] 6.2.3 When the influent level of the system's No. 10 buffer tank reaches its upper limit, the discharge will begin. Adjust the flow rate to 500 L / h for the membrane desalination system influent. Adjust the membrane desalination system parameters until the circulating desalination reaches 80% monovalent salt and 20% divalent salt. Monovalent brine enters the No. 11 buffer tank, and divalent brine enters the No. 12 buffer tank. The upper limit of the buffer tanks is 2.5 m. 3 Lower limit 0.5 m 3 .

[0299] 6.2.4 Salt composition design table 6.2.5.1 Monovalent Salts

[0300] 6.2.5.2 Divalent Salts

[0301] 6.2.6 The concentration of monovalent salts was 9.97%, and the concentration of divalent salts was 29.02%.

[0302] 6.2.7 RO Concentrated Water 8849.3L

[0303] 7. Evaporation and crystallization

[0304] 7.1 Evaporation and crystallization of monovalent salts

[0305] 7.1.1 When the influent level in the No. 11 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, feeding the No. 1 MVR evaporator. The feed flow rate is adjusted to 39.68 L / h, controlling the evaporation rate at 32.57 kg / h, and the evaporation rate of the No. 2 MVR evaporator is controlled at 3.96 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 1 crystallizer. The upper limit of the No. 1 crystallizer is 2.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0306] 7.1.2 When the upper limit of crystallizer #1 is reached, the discharge valve of crystallizer #1 is opened, and centrifuge #1 (GKH630-N) starts working, with the mother liquor entering buffer tank #12. When the centrifuge is full, the centrifuge backwash pump starts, rinsing the sodium chloride with saturated brine, and the rinse water enters buffer tank #12. After rinsing is complete, the centrifuge discharges the sodium chloride to the system's #1 sulfurized bed for drying and packaging, completing the sodium chloride recovery.

[0307] 7.1.3 Evaporation and crystallization of monovalent salts are completed.

[0308] 7.1.3.1 System 15 # Storage Tank Condensate: 694.03 kg

[0309] 64.14 kg of sodium chloride was recovered.

[0310] Use 35 kg of saturated brine for cleaning.

[0311] 7.1.3.2 Total Composition Table of Mother Liquor from Sodium Chloride Crystallization Centrifugation

[0312] 7.2 Evaporation and crystallization of divalent salts

[0313] 7.2.1 Design table of composition of divalent concentrated water before evaporation

[0314] 7.2.2 Separation of sodium chloride

[0315] 7.2.2.1 When the influent level in the system's No. 12 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, and feed is introduced into the No. 3 MVR evaporator. The feed flow rate is adjusted to 13.32 L / h, and the evaporation rate is controlled at 4.55 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 2 crystallizer. The upper limit of the No. 2 crystallizer is 2.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0316] 7.2.2.2 With the upper limit of crystallizer #2 set, the discharge valve of crystallizer #2 is opened, and centrifuge #2 (GK450-N) is started. The centrifuge discharges material into the sodium chloride storage silo (in industrial production, this sodium chloride is used to prepare saturated cleaning water and electrolyte; the excess is dissolved in monovalent brine for circulating evaporation and crystallization). The centrifuged mother liquor enters buffer tank #13. The upper limit of buffer tank #13 is 2.5m. 3 Lower limit 0.5m 3 .

[0317] 7.2.2.3 Separation complete.

[0318] 7.2.2.3.1 System 15 # Storage Tank Condensate: 86.4 kg

[0319] 15.65 kg of sodium chloride was recovered.

[0320] 7.2.2.3.2 Total Composition Table of Mother Liquor from Sodium Chloride Crystallization Centrifugation

[0321] 7.2.3 Separation of potassium chloride

[0322] 7.2.3.1 When the influent level in the system's No. 13 buffer tank reaches the lower limit, the evaporator system starts, and the MVR evaporator preheats. When the influent level reaches the upper limit, the discharge pump starts, and feed is introduced into the No. 4 MVR evaporator. The feed flow rate is adjusted to 10.17 L / h, and the evaporation rate is controlled at 1.75 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 3 crystallizer. The upper limit of the No. 3 crystallizer is 1.5 m. 3 Lower limit 0.5 m 3 Stir at 3 R / min.

[0323] 7.2.3.2 With the upper limit of crystallizer #3 set, the discharge valve of crystallizer #3 is opened, and centrifuge #3 (GK400-N) is started. The centrifuge discharges material into the potassium chloride #2 sulfurized bed for drying and packaging. The centrifuged mother liquor enters the buffer tank #14. The upper limit of buffer tank #14 is 0.3m. 3 Lower limit 0.1m 3 .

[0324] 7.2.3.3 Separation complete. 7.2.3.3.1

[0326] System 15 storage tank: 33.72 kg of condensate, 11.31 kg of potassium chloride recovered.

[0327] 7.2.3.3.2 Total Composition Table of Potassium Chloride Crystallization Centrifugation Mother Liquor

[0328] 7.2.4 Recovery of calcium chloride and magnesium chloride

[0329] 7.2.4.1 When the influent level in the system's No. 14 buffer tank reaches the lower limit, the evaporator system starts, and the multi-effect evaporator preheats. When the influent level reaches the upper limit, the forced circulation pump starts, and the multi-effect evaporator is fed and circulated. The circulation flow rate is adjusted to 7.6 L / h, and the evaporation rate is controlled at 4.46 kg / h. The evaporator condensate is collected in the No. 15 storage tank, and the evaporator concentrate enters the No. 4 crystallizer. The upper limit of the crystallizer is 0.3 m. 3 Lower limit 0.1m 3 Stir at 3 R / min.

[0330] 7.2.4.2 System #4 crystallizer upper limit, #4 crystallizer discharge valve open, #4 centrifuge (GK400-N) start, centrifuge discharges material into calcium chloride and magnesium chloride #3 sulfidation bed.

[0331] Drying and packaging.

[0332] 7.2.4.3 Separation complete.

[0333] 7.2.4.3.1 System 15 # Storage Tank Condensate: 84.76 kg

[0334] 119.15 kg of a mixed salt of calcium chloride and magnesium chloride was recovered.

[0335] 8. After processing 10 tons of water samples, a total of [amount] was recovered.

[0336] 8.1 Sodium chloride 79.79 kg,

[0337] 8.2 kJ of potassium chloride, 11.31 kg

[0338] 8.3 Calcium chloride and magnesium chloride 119.15 kg

[0339] 8.4 Evaporation condensate 9663.46 kg

[0340] 8.5% recovery rate: 2.22 kg crude oil

[0341] 8.6 Use 12.3L of 10% alkali and 84.2L of 10% barium chloride solution.

[0342] 8.7 Use 10.72L of 15% acid and 35kg of saturated brine.

[0343] 8.8 Solid waste 10.25kg

[0344] 8.9 Quality Table of Recycled Products

[0345] 8.9.1 Sodium chloride

[0346] 8.9.2 Potassium chloride

[0347] 8.9.3 Mixed salts of calcium chloride and magnesium chloride (snow-dissolving agent)

[0348] 8.9.4 Evaporation condensate

Claims

1. A method for treating wastewater from oil and gas field extraction, the method being: a. Removing impurities a1. Use a high-speed mixer to adjust the pH value of the gas field produced wastewater to 2-3. The acid used to adjust the pH value comes from the hydrochloric acid produced by the electrolysis of sodium chloride in the subsequent process. This reduces the solubility of oil in water, increases the solubility of inorganic salts in water, and fixes the upper limit of silicon solubility in water to ≤85mg / L. a2. A tubular / plate filter is used for separation and concentration. The separated oily water enters the next oil removal process. The separated reduced-volume slurry is pressed into a filter cake by a diaphragm filter press. The filter cake is washed and dried with the recovered water from the evaporation and crystallization process and treated as solid waste. The washing water is returned to the filter for re-filtration. This process does not add any chemical agents or additives (demulsifiers, surfactants, flocculants), only removing suspended solids and excess silica. The reduction process agitates the raw water, accelerating the demulsification of oil in the water and preparing it for the next oil removal process. b. Specialized membrane for oil removal b1, oil-water separation ultrafiltration membrane separates oil and water using a two-stage separation method. The final effluent contains ≤0.2mg / L of oil and ≤0.3mg / L of TSS. The membrane concentrate is returned to the previous impurity removal process, while the membrane permeate enters the next process for desiliconization and hardening removal. b2, Special membrane for oil removal b2.1 Adjust the pH of the dissolved air flotation product water to 6-8. The alkali used to adjust the pH comes from the mixed alkali produced by the electrolysis of sodium chloride and potassium chloride in the subsequent process. b2.2 Oil-water separation ultrafiltration membrane separates oil and water, using two-stage separation. The first stage has an oil-water separation rate of 90%, and the second stage has an oil-water separation rate of 99.8%. The separated oily concentrate is returned to the previous air flotation process, and the separated oil-free water enters the next evaporation pre-concentration process. c. Desiliconization to remove heavy metals and sulfate ions c1. Concentrate the oil-free water after oil removal using an MVR evaporation system, with the concentration ratio limited to the point where sodium chloride reaches saturation. The condensate from the evaporation is collected in a centralized drainage tank for testing and discharge, while the concentrated evaporation water enters the next process. c2. The pH value is adjusted from 2-3 to 11.1-11.5 using a high-speed mixer. The alkali used for pH adjustment comes from the alkali produced by the electrolysis of sodium chloride in the subsequent process. Heavy metals in the water form hydroxide precipitates, and silicon reacts with excess magnesium in the raw water to form magnesium silicate precipitates. c3. Solid-liquid separation is performed by combining tubular / plate-type volume reduction filters with filter presses. The clear water from the volume reduction filters enters the next degassing process, while the concentrated water and filter cake washing water from the filter press are returned to the front end of this process for circulation. After the filter cake is washed and dried, it is treated as waste or further recycled. The steam condensate generated during the drying process is collected in a centralized drainage pool for testing and discharge. c4. In this process, barium chloride is added at the theoretical value of 0.95-0.98 during sulfate removal. Barium does not leave residues in the water; it is merely a replacement of chloride ions with sulfate ions, resulting in a decrease in the variety of components in the water, not an increase. Compared with other sulfate removal methods, the cost is lower, and scale inhibitors and defoamers are not required in the subsequent membrane and evaporation crystallization processes. Under alkaline conditions, silicon combines with magnesium in the raw water to form insoluble magnesium silicate, while other heavy metals form insoluble hydroxides or dehydrate to precipitate oxides, achieving desiliconization and hardness removal. The residual silicon and heavy metals in the water meet the requirements of the subsequent salt separation process. d. Degassing to remove methane, methanol, inorganic ammonia, and small-molecule organic matter from wastewater. d1. After the wastewater is treated for impurities, oil, silicon and hardness, in addition to sodium, potassium, calcium, magnesium and trace amounts of silicon and heavy metals, there are also small molecule organic matter and inorganic ammonia in the water. The organic matter is mainly methane and methanol, which need to be removed before salt separation. d2. A vacuum microwave degassing tower is used to atomize alkaline mixed salt water under negative pressure, and the droplets are heated by microwaves. Methane and inorganic ammonia are desorbed under negative pressure, and methanol in the droplets is preferentially desorbed by microwave absorption before water. Other small molecule organic compounds are also desorbed during the heating process, leaving the water body and entering the gas collection tank to complete the degassing. This process is more energy-efficient than traditional steam stripping, does not require the addition of steam and does not increase the subsequent evaporation, and has a large amount of small molecule organic compounds desorbed. e. Heterogeneous ozone catalytic oxidation e1, the degassed alkaline mixed brine is further treated with an ozone catalytic oxidation device to remove residual organic matter and ammonia; e2, the oxidized water is filtered by ultrafiltration, the ultrafiltration concentrate is returned to the desiliconization and heavy metal removal process, and the ultrafiltration permeate enters the next process for salt separation. f, Membrane separation, membrane separation of monovalent and divalent salts. f1, Adjust the pH of the mixed brine to 6.

5. The acid used to adjust the pH value comes from the hydrochloric acid produced by the electrolysis of sodium chloride in the subsequent process; f2 uses a (100 Da) nanofiltration membrane to separate salts, with a monovalent salt concentration ≥10% and a divalent salt concentration ≤30%; g Evaporation and Crystallization g1. Evaporation and crystallization separation of monovalent salts g1.1 Evaporate monovalent salts to precipitate sodium chloride using an MVR single crystallization evaporation system until potassium chloride, calcium chloride, and magnesium chloride are co-saturated; g1.2 Recover sodium chloride using a horizontal siphon scraper centrifuge. The centrifuge drum is equipped with a saturated brine spray pipe, which directly sprays the filter cake at full speed to wash away trace amounts of organic mother liquor and ammonium impurities carried on the crystal surface, thereby improving the salt quality. The saturated brine comes from the sodium chloride recovered by evaporation of divalent brine and the monovalent brine separated by the membrane. g1.

3. Dry sodium chloride in a sulfide bed until the moisture content is within acceptable limits; g1.4 The recovered sodium chloride meets the industry standard T / ZGZS0302-2023; g1.

5. The final mother liquor from the evaporation of monovalent salts enters the divalent salt evaporation and crystallization process for further evaporation and crystallization. g2, evaporation and crystallization separation of divalent salt water g2.1 Separate sodium chloride by evaporating divalent salts using an MVR evaporator until potassium chloride, calcium chloride, and magnesium chloride are co-saturated; g2.1.1 Sodium chloride is recovered by using a horizontal siphon scraper centrifuge. The centrifuge drum is equipped with a saturated brine spray pipe, which directly sprays the filter cake at full speed to wash away trace amounts of organic mother liquor and ammonium impurities on the crystal surface, thereby improving the salt quality. The saturated brine comes from sodium chloride recovered by evaporation of divalent brine and monovalent brine separated by membrane. g2.1.

2. Part of the precipitated sodium chloride is mixed with the monovalent brine of the membrane to prepare a sodium chloride saturated cleaning solution, and part of it is mixed with the monovalent brine of the membrane to prepare an electrolyte; the remaining part is returned to the monovalent brine of the membrane separation for recrystallization, which only improves the purity of sodium chloride separated by divalent salt, without changing the evaporation rate of monovalent brine. g2.

2. Continue to use an MVR evaporator to crystallize and separate potassium chloride from the mother liquor after separating sodium chloride, until calcium chloride and magnesium chloride are co-saturated; g2.2.

1. Recover potassium chloride using a horizontal centrifuge and dry it with a sulfide bed until the moisture content is within acceptable limits; g2.2.2 The recovered potassium chloride meets the Class II standard of the National Industrial Salt (GB6549-2011); g2.3 The final mother liquor is circulated and evaporated using a multi-effect evaporator to crystallize a mixed salt of calcium chloride and magnesium chloride; g2.3.1 Use a horizontal centrifuge to recover the mixed salt of calcium chloride and magnesium chloride, and dry it with a sulfide bed until the moisture content is qualified; g2.3.2, The production of de-icing agents shall meet the national standard GB / T 23851-2017; h, Sodium chloride electrolysis The electrolyte for the h1 membrane was prepared by evaporating and crystallizing monovalent salt water and divalent salt sodium chloride. H2 brine is refined by using chelating resin to remove cations, meeting the requirements for bipolar membrane electrolysis feed water. h2.1 After the resin is saturated, it is regenerated with electrolytic hydrochloric acid, and the regenerated water is returned to the divalent salt for evaporation and crystallization; h3 performs bipolar membrane electrolysis on qualified brine. This electrolysis process directly produces alkali and hydrochloric acid without producing chlorine gas. The produced alkali and hydrochloric acid are used to adjust the pH value in each process, and the amount of electrolysis is based on the requirements of each process. After the pH value is adjusted, sodium hydroxide and hydrochloric acid are reduced back to sodium chloride. They only play a conditional control role and do not actually participate in the recombination of other elements to generate new substances, thus not changing the total amount of components in the gas field water treatment process.

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