Method for removing oil phenol and ammonia from high-concentration phenol-containing coal chemical wastewater

Abstract

The present application belongs to the technical field of wastewater treatment, and discloses a method for removing oil phenol and ammonia from high-concentration phenol-containing coal chemical wastewater. The method comprises the following steps: extracting oil from the coal chemical wastewater by using n-butyl ether to obtain an oil-removed raffinate phase and an oil-containing extract phase; simultaneously performing deacidification and deamination treatment on the oil-removed raffinate phase; and then extracting phenol from the wastewater after deacidification and deamination by using n-butyl ether to obtain a phenol-containing extract phase and wastewater from which oil phenol and ammonia have been removed; wherein the oil-containing extract phase and the phenol-containing extract phase are subjected to rectification treatment to recover n-butyl ether for reuse. The present application uses n-butyl ether as an extractant, which has good extraction efficiency for polyphenol. The present application uses an oil-removing extraction tower, a deacidification and deamination tower, a phenol-removing extraction tower and a solvent stripping tower to remove oil, ammonia and phenol, greatly simplifying the process of removing oil, ammonia and phenol. In the present application, alkali liquor is added to a position below the side line tapping position and above the tower kettle of the deacidification and deamination tower, which can better remove fixed ammonium.

Description

TECHNICAL FIELD

[0001] The present application belongs to the technical field of wastewater treatment, and particularly relates to a method for removing oil, phenol and ammonia from high-concentration phenol-containing coal chemical wastewater. BACKGROUND

[0002] In the process of gasification or dry distillation, low-rank coal produces high-concentration organic and inorganic pollutants such as oil, phenol, ammonia and other coal pollutants, and a large amount of high-concentration coal chemical wastewater containing phenol is produced. If these wastewaters are directly discharged without treatment, they will cause great pollution to the natural environment. In particular, the oil, ammonia and phenol substances have relatively high concentrations and strong toxicity. Therefore, removing the oil, ammonia and phenol substances in high-concentration coal chemical wastewater is a key link in the comprehensive treatment of high-concentration coal chemical wastewater containing phenol.

[0003] Currently, the mainstream method for removing phenol and ammonia from high-concentration coal chemical wastewater containing phenol is to first remove acid and ammonia from the phenol-containing wastewater after oil and dust removal, then extract and remove phenol, and then use a phenol tower and a water tower to recover and recycle the extractant in the extraction phase and the raffinate phase, respectively, to obtain crude phenol and wastewater after removal of phenol and ammonia. The main extractants are methyl isobutyl ketone and diisopropyl ether. For Lurgi gasification wastewater, due to the low content of polyphenol, diisopropyl ether can reduce the COD to 5000, and methyl isobutyl ketone can reduce the COD to below 3000 or even lower. For high-concentration polyphenol-containing coke wastewater, diisopropyl ether is completely unsuitable, and methyl isobutyl ketone needs to be used as an extractant to reduce the wastewater COD from 40,000 to 55,000 or below. The above processes all have problems such as high oil content entering the phenol and ammonia removal unit, causing the plate of the acid and ammonia removal tower to be blocked, high oil content of the ammonia gas discharged by the three-stage condensation, and affecting the efficiency of the extraction and removal of phenol. Chinese Patent Publication CN106315906A proposes to add an extraction reactor before acid and ammonia removal, and then use five towers including an acid rectification tower, an ammonia rectification tower, a phenol oil extraction tower, a phenol back extraction tower and an extractant purification tower to remove phenol and ammonia; a mixed extractant is used, in which the solubility of 3-heptanone and propylene glycol butyl ether in water at 20°C is 3.3 g / L and 6.4%, respectively, i.e. there is a loss of nearly 3.3 kg and 64 kg per ton of water treated, and the COD contribution in water is also large, so a water tower is usually needed to recover and utilize the extractant; in addition, the process has problems such as the position of adding alkali for ammonia removal, back extraction after adding alkali, and the like, and the process is complicated. SUMMARY

[0004] In order to overcome the shortcomings and deficiencies of the prior art, the primary purpose of the present application is to provide a method for removing oil, phenol and ammonia from high-concentration phenol-containing coal chemical wastewater.

[0005] The method of the present application removes oil phenol and ammonia by four towers, and the process is simple and clear. The same extractant with low solubility in water is used in the oil removal extraction tower and the phenol removal extraction tower, which reduces the interference of the extractant on the deacidification and deamination, and saves the investment and daily operation cost of the water tower. The oil removal improves the efficiency of the subsequent deacidification, deamination and extraction of phenol.

[0006] The object of the present application is achieved by the following scheme:

[0007] A method for removing oil, phenol and ammonia from high-concentration coal chemical wastewater, comprising the following steps: extracting oil from the coal chemical wastewater by using n-butyl ether to obtain an oil-removal raffinate phase and an oil-containing extraction phase; simultaneously treating the oil-removal raffinate phase to remove acid and ammonia; then extracting phenol from the wastewater after the removal of acid and ammonia by using n-butyl ether to obtain a phenol-containing extraction phase and wastewater from which oil, phenol and ammonia have been removed; and recycling the n-butyl ether in the oil-containing extraction phase and the phenol-containing extraction phase by rectification.

[0008] The coal chemical wastewater has a COD content of 22500-51200 mg / L, an oil content of 2140-3130 mg / L, a total phenol content of 6800-15100 mg / L, and an ammonia content of 5600-9770 mg / L.

[0009] Preferably, the coal chemical wastewater has a COD content of 24500-48600 mg / L, an oil content of 2430-3000 mg / L, a total phenol content of 6800-14400 mg / L, and an ammonia content of 5600-8800 mg / L.

[0010] A method for removing oil, phenol and ammonia from high-concentration coal chemical wastewater, comprising the following steps:

[0011] (1) Oil removal extraction: n-butyl ether and the coal chemical wastewater enter an oil removal extraction tower for extraction to obtain an oil-containing extraction phase and an oil-removal raffinate phase;

[0012] (2) Deacidification and deamination: the oil-removal raffinate phase is divided into two paths and enters a deacidification and deamination tower, the cold feed enters the upper part of the packing section of the deacidification and deamination tower from the top, the hot feed enters the first tray below the packing section, and alkali liquor is added to the deacidification and deamination tower; acidic gas is obtained at the top, ammonia gas is obtained in the middle part, and deacidification and deamination wastewater is obtained at the bottom;

[0013] (3) Extraction and phenol removal: n-butyl ether and the deacidification and deamination wastewater enter a phenol removal extraction tower for extraction to obtain a phenol-containing extraction phase and wastewater from which oil, phenol and ammonia have been removed;

[0014] (4) Solvent stripping: the oil-containing extraction phase of step (1) and the phenol-containing extraction phase of step (3) enter a solvent stripping tower, n-butyl ether is obtained at the top and is sent to a solvent circulation tank for recycling after heat exchange and cooling, and crude phenol and tar are obtained at the bottom.

[0015] The extraction in step (1) and step (3) is countercurrent extraction.

[0016] The extraction in step (1) is carried out at a temperature of 30-90℃; the extraction stage number of the oil-removing extraction column is 2-8; the theoretical tray number is the same as the extraction stage number.

[0017] The volume ratio of n-butyl ether to coal chemical wastewater in step (1) is 1:9-1:3.

[0018] In step (2), the addition of lye is specifically adding sodium hydroxide solution into the deacidification and deamination column from the position above the column bottom and below the side line extraction position, preferably from the 5th-7th tray below the side line extraction position; the amount of lye used is sufficient to completely convert the fixed ammonium in the wastewater into ammonia water.

[0019] The fixed ammonium is at least one of ammonium chloride, ammonium sulfate and ammonium nitrate.

[0020] The deacidification and deamination column in step (2) is a composite column combining a packed column and a plate column.

[0021] In step (2), the cold feed inlet temperature is 30-60℃, the hot feed inlet temperature is 130-170℃, and the column bottom pressure is 0.3-0.7MPa.

[0022] In step (2), the ammonia gas obtained from the middle part of the column is specifically crude ammonia gas extracted from the side line of the ammonia gas enrichment zone, then enters a three-stage fractional condensation, and is flashed out after being cooled and decompressed, and the side line extraction rate is 8%-15%.

[0023] The three-stage fractional condensation includes a first-stage flash tank with a flash pressure of 0.25-0.45MPa and a flash temperature of 120-140℃, a second-stage flash tank with a flash pressure of 0.2-0.4MPa and a flash temperature of 70-110℃, and a third-stage flash tank with a flash pressure of 0.10-0.25MPa and a flash temperature of 30-50℃.

[0024] The extraction in step (3) is carried out at a temperature of 30-90℃, and the extraction stage number of the dephenolization extraction column is 3-10; the theoretical tray number is the same as the extraction stage number.

[0025] The volume ratio of n-butyl ether to deacidification and deamination wastewater in step (3) is 1:4-1:1.

[0026] In step (4), the overhead temperature of the solvent stripping column is 125-170℃, the column bottom temperature is 195-240℃, the overhead pressure is 0.1-0.2MPa, and the theoretical tray number is 15-35.

[0027] In step (4), the phenol-containing extraction phase and the oil-containing extraction phase are heated and warmed after being combined with the overhead discharge of the solvent stripping column.

[0028] The present application has the following advantages and beneficial effects relative to the prior art:

[0029] (1) The present application selects n-butyl ether as an extractant, which has good extraction efficiency on polyphenol but very low solubility in water, to perform oil removal extraction before phenol and ammonia removal, thereby reducing the oil content of waste water entering the phenol and ammonia removal process, greatly slowing down the plugging of the tray in the ammonia removal device, simplifying ammonia refining and obtaining high-quality ammonia water, and being beneficial to subsequent extraction and phenol removal.

[0030] (2) The present application uses the same extractant for oil removal extraction and phenol removal extraction, simplifies the process, and can directly use the same rectification tower for solvent stripping to recover n-butyl ether.

[0031] (3) The present application uses n-butyl ether with low solubility in water, which does not dissolve new extractants before acid and ammonia removal, thereby increasing the complexity of the acid and ammonia removal tower and not causing any interference to the removal of acid gas and ammonia.

[0032] (4) The present application uses a four-tower process of an oil removal extraction tower, an acid and ammonia removal tower, a phenol removal extraction tower, and a solvent stripping tower to remove oil, ammonia, and phenol, greatly simplifying the oil, phenol, and ammonia removal process.

[0033] (5) The present application adds alkali liquor below the side line extraction position and above the tower kettle of the acid and ammonia removal tower, which can better remove fixed ammonium: Ammonia has two main forms in aqueous solution, one is free ammonium, such as ammonium carbonate, ammonium bicarbonate, ammonium sulfide, ammonium hydrosulfide, ammonia water, etc., which usually decomposes at 100°C or higher; the other is fixed ammonium, such as ammonium chloride, ammonium sulfate, ammonium nitrate, etc., which hardly decomposes in water below 350°C. Fixed ammonium usually needs to be converted into ammonia water by adding alkali to remove ammonia. From the perspective of removing fixed ammonium, the closer the alkali addition position to the upper part of the acid and ammonia removal tower, the more time the fixed ammonium has to be converted into free ammonia. However, the closer to the upper part of the tower, the higher the concentration of carbonate ions, bicarbonate ions, and sulfide ions in the liquid phase, and the higher the concentration of components such as sodium carbonate, sodium bicarbonate, and sodium hydrosulfide. Hydrogen sulfide and carbon dioxide, etc. are fixed, which reduces the acid gas removal efficiency of the acid and ammonia removal tower. Therefore, the present application selects to add alkali liquor below the side line extraction position and above the tower kettle, which can better remove fixed ammonium.

[0034] (6) The present application does not use alkali liquor back extraction when recovering the solvent, avoiding the increase of the pH value of the extraction by alkali, which is not conducive to the extraction of oil and phenol removal. BRIEF DESCRIPTION OF DRAWINGS

[0035] Figure 1 The present application is a schematic diagram of the high-concentration phenol-containing coal chemical wastewater oil, phenol, and ammonia removal method.

[0036] The numbers in the figure are explained as follows: 1, coal chemical wastewater tank, 2, oil removal extraction column, 3, extraction agent, 4, lye, 5, cooler, 6, deacidification and deamination column, 7, acid gas, 8, stream heat exchanger, 9, first-stage flash tank, 10, condenser, 11, second-stage flash tank, 12, crude ammonia condenser, 13, third-stage flash tank, 14, ammonia gas, 15, cooler, 16, circulating wastewater, 17, reboiler, 18, stream heat exchanger, 19, cooler, 20, dephenol extraction column, 21, stream heat exchanger, 22, cooler, 23, fresh extraction agent, 24, wastewater from which oil, phenol and ammonia are removed, 25, extraction agent circulating tank, 26, solvent stripping column, 27, cooler, 28, reboiler, 29, cooler, 30, crude phenol or tar. DETAILED DESCRIPTION

[0037] The present application will be described in further detail below with reference to Examples and the accompanying drawings, but the embodiments of the present application are not limited thereto. In the Examples, the specific conditions not mentioned are performed according to the conventional conditions or the conditions recommended by the manufacturers. The reagents or instruments used are not mentioned by the manufacturers, and are all conventional products that can be purchased on the market.

[0038] The reagents used in the Examples are all conventional products that can be purchased on the market unless otherwise specified.

[0039] Example 1

[0040] 85 tons / hour of Lurgi gasification wastewater, the COD content of which is 24500 mg / L, the oil content is 3000 mg / L, the total phenol content is 6800 mg / L, and the ammonia content is 8800 mg / L.

[0041] The n-butyl ether and the high-concentration phenol-containing coal chemical wastewater enter the oil removal extraction column for countercurrent extraction, the extraction temperature is 60°C, the extraction stage (theoretical tray number) is 5, and the volume ratio of the n-butyl ether to the wastewater is 1:5; the oil-containing extraction phase is discharged from the top of the column and sent to the solvent stripping column, and the oil removal extraction raffinate is discharged from the column bottom and sent to the deacidification and deamination column.

[0042] The deoiled raffinate phase from the deoiling extraction column is divided into two streams, one cold and one hot, which are fed into a deacidification and deamination column, which is a composite column combining a packed column and a plate column. One stream is cooled to a temperature of 47°C by a cooler and then fed into the upper part of the packed column from the top of the deacidification and deamination column, and the other stream is heated to a temperature of 136°C by heat exchange and then fed into the first plate below the packed section of the deacidification and deamination column, and sodium hydroxide solution is added from the 6th plate below the side line extraction position. The cold feed and the hot feed are combined and heat exchanged with steam to strip out ammonia gas and acidic gas. The column top pressure is 0.4 MPa, and the ammonia gas is enriched in the middle part of the column, and the side line steam is extracted and enters a three-stage flash evaporation, with a side line extraction rate of 9%. The three-stage flash evaporation includes: a first-stage flash tank with a flash pressure of 0.4 MPa and a flash temperature of 130°C; a second-stage flash tank with a flash pressure of 0.3 MPa and a flash temperature of 95°C; and a third-stage flash tank with a flash pressure of 0.25 MPa and a flash temperature of 40°C, after which the ammonia gas is flashed out after being cooled and depressurized, and the liquid phase of the flash tanks is combined and returned to the wastewater tank. The deacidification and deamination wastewater from the bottom of the deacidification and deamination column is heat exchanged and cooled and then fed into the dephenol extraction column.

[0043] The n-butyl ether from the solvent circulation tank is countercurrently extracted with the deacidification and deamination wastewater from the bottom of the deacidification and deamination column to remove phenol, with an extraction temperature of 60°C, an extraction stage (theoretical plate number) of 4, and a volume ratio of n-butyl ether to deacidification and deamination wastewater of 1:1.3, to obtain a phenol-containing extraction phase and an oil-phenol-ammonia-removed wastewater, which is fed into a subsequent biochemical treatment unit.

[0044] Solvent stripping: the extraction phases from the deoiling extraction column and the dephenol extraction column are combined and heat exchanged with the overhead material of the solvent stripping column to be heated and then fed into the solvent stripping column, which uses rectification to recover the n-butyl ether, with a column top temperature of 139°C, a column bottom temperature of 218°C, a column top pressure of 0.1 MPa, and a theoretical plate number of 21. The n-butyl ether is distilled out of the column top and is heat exchanged and cooled after being extracted to be recycled into the solvent circulation tank, and the column bottom obtains crude phenol and tar.

[0045] The dephenol extraction column water (i.e., the oil-phenol-ammonia-removed wastewater) has a COD content of 2450 mg / L, an oil content of 50 mg / L, a total phenol content of 245 mg / L, and an ammonia content of 178 mg / L.

[0046] Example 2:

[0047] 85 tons / hour of Lurgi gasification wastewater has a COD content of 24500 mg / L, an oil content of 3000 mg / L, a total phenol content of 6800 mg / L, and an ammonia content of 8800 mg / L.

[0048] The n-butyl ether and the high-concentration phenol-containing coal chemical wastewater enter a countercurrent extraction tower for oil removal extraction, the extraction temperature is 40°C, the extraction stage (theoretical tray number) is 8, the volume ratio of the n-butyl ether to the wastewater is 1:6. The oil-containing extraction phase is discharged from the top of the tower and is sent to a solvent stripping tower, and the oil-removal extraction raffinate phase is discharged from the tower bottom and is sent to an acid and ammonia removal tower.

[0049] The oil-removal extraction raffinate phase discharged from the oil removal extraction tower is divided into two streams, which are cold and hot, respectively, and enter the acid and ammonia removal tower (a combined tower of a packed tower and a plate tower). One stream is cooled to 40°C by a cooler and then enters the upper part of the packed tower from the top of the acid and ammonia removal tower, and the other stream is heated to 145°C by heat exchange and then enters the first tray below the packed section of the acid and ammonia removal tower, and sodium hydroxide solution is added from the fifth tray below the side line extraction position; the cold feed and the hot feed are combined and heat-exchanged with steam to strip out ammonia gas and acid gas. The tower top pressure is 0.6 MPa, and the acid gas is flashed out after the tower top steam is cooled; the ammonia gas is enriched in the middle part of the tower, and the side line extracted steam enters a three-stage flash evaporation, and the side line extraction rate is 13%. The three-stage flash evaporation includes: a first-stage flash tank, the flash pressure is 0.35 MPa, and the flash temperature is 125°C; a second-stage flash tank, the flash pressure is 0.25 MPa, and the flash temperature is 70°C; a third-stage flash tank, the flash pressure is 0.15 MPa, and the flash temperature is 35°C, and the ammonia gas is flashed out after being cooled and depressurized, and the liquid phase of the flash tanks is combined and then returned to the wastewater tank; the acid and ammonia removal wastewater is obtained from the tower bottom, and is sent to a phenol extraction tower after being heat-exchanged and cooled.

[0050] The n-butyl ether from the solvent circulation tank and the acid and ammonia removal wastewater from the tower bottom of the acid and ammonia removal tower are countercurrently extracted to remove phenol, the extraction temperature is 40°C, the extraction stage (theoretical tray number) is 5, the volume ratio of the n-butyl ether to the acid and ammonia removal wastewater is 1:1.67, and the phenol-containing extraction phase and the oil-phenol-ammonia-removed wastewater are obtained, and the oil-phenol-ammonia-removed wastewater enters a subsequent biochemical treatment unit.

[0051] Solvent stripping: the extraction phases from the oil removal extraction tower and the phenol extraction tower are combined, heat-exchanged and heated with the tower top discharge of the solvent stripping tower, and then enter the solvent stripping tower, and the n-butyl ether is recovered by rectification, the tower top temperature is 125°C, the tower bottom temperature is 195°C, the tower top pressure is 0.11 MPa, and the theoretical tray number is 35. The n-butyl ether is distilled out from the tower top, is cooled and heat-exchanged after being extracted, and is sent to the solvent circulation tank for recycling, and the crude phenol and tar are obtained from the tower bottom.

[0052] The COD content of the water discharged from the phenol extraction tower (i.e., the oil-phenol-ammonia-removed wastewater) is 2231 mg / L, the oil content is 87 mg / L, the total phenol content is 216 mg / L, and the ammonia content is 135 mg / L.

[0053] Example 3:

[0054] 85 tons / hour of Lurgi gasification wastewater, COD content of 24500 mg / L, oil content of 3000 mg / L, total phenol content of 6800 mg / L, ammonia content of 8800 mg / L.

[0055] n-Butyl ether and high concentration phenolic coal chemical wastewater into the oil extraction tower countercurrent extraction, extraction temperature is 50 ℃, extraction stage (theoretical plate number) is 6, n-butyl ether and wastewater volume ratio is 1:4; oil-containing extraction phase from the tower top, sent to the solvent stripping tower, oil extraction raffinate phase from the tower bottom, sent to the deacidification and deamination tower.

[0056] The oil removal raffinate phase from the oil removal extraction tower is divided into cold and hot streams, which are respectively introduced into the deacidification and deamination tower (a combined tower of packed tower and plate tower). One stream is cooled to 45 ℃ by a cooler and then introduced into the upper part of the packed tower from the top of the deacidification and deamination tower, and the other stream is heated to 130 ℃ by heat exchange and then introduced into the first plate below the packed section of the deacidification and deamination tower, and sodium hydroxide solution is added from the 6th plate below the side line extraction position; the cold feed and the hot feed are combined and heat exchanged with steam to strip out ammonia gas and acid gas. The tower top pressure is 0.45 MPa, and the acid gas is flashed out after the tower top steam is cooled. At the ammonia gas enrichment position in the middle part of the tower, the side line extracted steam is introduced into a three-stage flash evaporation, and the side line extraction rate is 11%. The three-stage flash evaporation includes: a first-stage flash tank with a flash pressure of 0.45 MPa and a flash temperature of 130 ℃; a second-stage flash tank with a flash pressure of 0.4 MPa and a flash temperature of 85 ℃; and a third-stage flash tank with a flash pressure of 0.2 MPa and a flash temperature of 30 ℃, and ammonia gas is flashed out after being cooled and depressurized, and the liquid phase of the flash tanks is combined and then returned to the wastewater tank; the deacidification and deamination wastewater from the tower bottom is sent to the dephenol extraction tower after being heat exchanged and cooled.

[0057] The n-butyl ether from the solvent circulation tank and the deacidification and deamination wastewater from the tower bottom of the deacidification and deamination tower are countercurrently extracted to remove phenol, the extraction temperature is 80 ℃, the extraction stage (theoretical plate number) is 8, the volume ratio of n-butyl ether to deacidification and deamination wastewater is 1:1, and the phenol-containing extraction phase and the oil-phenol-ammonia-removed wastewater are obtained, and the oil-phenol-ammonia-removed wastewater is introduced into a subsequent biochemical treatment unit.

[0058] Solvent stripping: the extraction phases from the oil extraction tower and the dephenol extraction tower are combined and then introduced into the solvent stripping tower after being heat exchanged and heated with the overhead material of the solvent stripping tower, n-butyl ether is recovered by rectification, the tower top temperature is 143 ℃, the tower bottom temperature is 223 ℃, the tower top pressure is 0.12 MPa, and the theoretical plate number is 24. N-butyl ether is distilled out from the tower top, cooled and sent to the solvent circulation tank for recycling after being heat exchanged, and crude phenol and tar are obtained from the tower bottom.

[0059] The COD content of the water from the dephenol extraction column (i.e., the wastewater from which oil phenol ammonia has been removed) is 2766 mg / L, the oil content is 43 mg / L, the total phenol content is 268 mg / L, and the ammonia content is 166 mg / L.

[0060] Example 4:

[0061] The 45 tons / hour of semicoke wastewater has a COD content of 48600 mg / L, an oil content of 2430 mg / L, a total phenol content of 14400 mg / L, and an ammonia content of 5600 mg / L.

[0062] The n-butyl ether and the high-concentration phenol-containing coal chemical wastewater enter the deoiling extraction column for countercurrent extraction, the extraction temperature is 30°C, the extraction stage (theoretical tray number) is 7, and the volume ratio of the n-butyl ether to the wastewater is 1:3; the oil-containing extraction phase is discharged from the top of the column and sent to the solvent stripping column, and the deoiling raffinate phase is discharged from the column bottom and sent to the deacidification and deamination column.

[0063] The deoiling raffinate phase from the deoiling extraction column is divided into two streams, which enter the deacidification and deamination column (a composite column combining a packed column and a plate column) respectively. One stream is cooled to a temperature of 30°C by a cooler and then enters the upper part of the packed column from the top of the deacidification and deamination column, and the other stream is heated to a temperature of 140°C by heat exchange and then enters the first tray below the packed section of the deacidification and deamination column, and sodium hydroxide solution is added from the 7th tray below the side line extraction position; the cold feed and the hot feed are combined and heat-exchanged with steam to strip out ammonia gas and acidic gas. The column top pressure is 0.3 MPa, and the ammonia gas and acidic gas are flashed out after the steam is cooled. At the ammonia gas enrichment position in the middle part of the column, the steam is extracted from the side line and enters a three-stage flash evaporation, and the side line extraction rate is 15%. The three-stage flash evaporation includes: a first-stage flash tank with a flash pressure of 0.42 MPa and a flash temperature of 120°C; a second-stage flash tank with a flash pressure of 0.27 MPa and a flash temperature of 80°C; and a third-stage flash tank with a flash pressure of 0.1 MPa and a flash temperature of 45°C, and the ammonia gas is flashed out after being cooled and depressurized, and the liquid phase of the flash tanks is combined and returned to the wastewater tank; the deacidification and deamination wastewater from the column bottom is sent to the dephenol extraction column after being heat-exchanged and cooled.

[0064] The n-butyl ether from the solvent circulation tank and the deacidification and deamination wastewater from the column bottom of the deacidification and deamination column are countercurrently extracted to remove phenol, the extraction temperature is 30°C, the extraction stage (theoretical tray number) is 10, and the volume ratio of the n-butyl ether to the deacidification and deamination wastewater is 1:2, and the phenol-containing extraction phase and the wastewater from which oil phenol ammonia has been removed are obtained, and the wastewater from which oil phenol ammonia has been removed is sent to a subsequent biochemical treatment unit.

[0065] Solvent stripping: the extraction phase from the de-oil extraction column and the de-phenol extraction column is combined, warmed by heat exchange with the overhead discharge of the solvent stripping column, and then fed to the solvent stripping column. The n-butyl ether is recovered by rectification. The overhead temperature is 154°C, the column bottom temperature is 229°C, the overhead pressure is 0.15 MPa, and the number of theoretical plates is 30. The n-butyl ether is distilled overhead and, after being cooled by heat exchange, is fed to the solvent circulation tank for recycling. The column bottom is obtained as crude phenol and tar.

[0066] The water from the de-phenol extraction column (i.e., the waste water from which the oil phenol and ammonia have been removed) has a COD content of 3274 mg / L, an oil content of 40 mg / L, a total phenol content of 256 mg / L, a polyphenol content of 245 mg / L, and an ammonia content of 182 mg / L.

[0067] Example 5:

[0068] 45 tons / hour of semi-coke waste water has a COD content of 48600 mg / L, an oil content of 2430 mg / L, a total phenol content of 14400 mg / L, and an ammonia content of 5600 mg / L.

[0069] The n-butyl ether and the high-concentration phenol-containing coal chemical waste water are fed to the de-oil extraction column for countercurrent extraction. The extraction temperature is 75°C, the number of extraction stages (theoretical plates) is 2, and the volume ratio of n-butyl ether to waste water is 1:7. The oil-containing extraction phase is discharged from the top of the column and fed to the solvent stripping column. The de-oil raffinate is discharged from the column bottom and fed to the de-acidification and de-ammoniation column.

[0070] The de-oil raffinate from the de-oil extraction column is divided into two streams, a cold stream and a hot stream, which are fed to the de-acidification and de-ammoniation column, which is a composite column combining a packed column and a plate column. One stream is cooled to a temperature of 56°C by a cooler and fed to the upper part of the packed column from the top of the de-acidification and de-ammoniation column. The other stream is warmed by heat exchange to a temperature of 170°C and fed to the first plate below the packed section of the de-acidification and de-ammoniation column. Sodium hydroxide solution is fed from the fifth plate below the side line extraction position. The cold feed and the hot feed are combined and heat-exchanged with steam to strip out the ammonia and acid gases. The overhead pressure is 0.62 MPa, and the acid gases are flashed out after being cooled by steam. The ammonia is enriched in the middle part of the column, and the steam is extracted from the side line and fed to a three-stage flash evaporation. The side line extraction rate is 8%. The three-stage flash evaporation includes: a first-stage flash tank with a flash pressure of 0.38 MPa and a flash temperature of 135°C; a second-stage flash tank with a flash pressure of 0.32 MPa and a flash temperature of 80°C; and a third-stage flash tank with a flash pressure of 0.24 MPa and a flash temperature of 48°C. The ammonia is flashed out after being cooled and depressurized, and the liquid phase from the flash tanks is combined and returned to the waste water tank. The de-acidification and de-ammoniation column bottom is obtained as de-acidification and de-ammoniation waste water, which is cooled by heat exchange and fed to the de-phenol extraction column.

[0071] The n-butyl ether from the solvent circulation tank is countercurrently extracted with the deacidification and deamination wastewater from the deacidification and deamination column, the extraction temperature is 56°C, the extraction stage (theoretical tray number) is 6, the volume ratio of n-butyl ether to deacidification and deamination wastewater is 1:3, and the phenol-containing extraction phase and the oil-phenol-ammonia-removed wastewater are obtained, and the oil-phenol-ammonia-removed wastewater enters the subsequent biochemical treatment unit.

[0072] Solvent stripping: the extraction phases from the oil removal extraction column and the phenol removal extraction column are combined, heated and then enter the solvent stripping column, the n-butyl ether is recovered by rectification, the top temperature is 166°C, the bottom temperature is 240°C, the top pressure is 0.18Mpa, and the theoretical tray number is 15. The n-butyl ether is distilled from the top of the column, cooled by heat exchange, and then sent to the solvent circulation tank for recycling, and the crude phenol and tar are obtained from the bottom of the column.

[0073] The COD content of the water from the phenol removal extraction column (i.e. the oil-phenol-ammonia-removed wastewater) is 4768mg / L, the oil content is 76mg / L, the total phenol content is 342mg / L, and the ammonia content is 148mg / L.

[0074] Example 6:

[0075] 45 tons / hour of semi-coke wastewater, the COD content is 48600mg / L, the oil content is 2430mg / L, the total phenol content is 14400mg / L, and the ammonia content is 5600mg / L.

[0076] The n-butyl ether and the high-concentration phenol-containing coal chemical wastewater enter the oil removal extraction column for countercurrent extraction, the extraction temperature is 90°C, the extraction stage (theoretical tray number) is 5, the volume ratio of n-butyl ether to wastewater is 1:9; the oil-containing extraction phase is discharged from the top of the column and sent to the solvent stripping column, and the oil-removal raffinate is discharged from the bottom of the column and sent to the deacidification and deamination column.

[0077] The deoiled raffinate phase from the deoiling extraction column is divided into two streams, one cold and one hot, which are fed into the deacidification and deamination column, which is a combined column of packed tower and plate tower. One stream is cooled to 60°C by a cooler and then fed into the upper part of the packed tower from the top of the deacidification and deamination column, and the other stream is heated to 160°C by heat exchange and then fed into the first plate below the packed section of the deacidification and deamination column, and sodium hydroxide solution is added from the 7th plate below the side draw point. The cold feed and the hot feed are combined and heat exchanged with steam to strip out ammonia and acid gas. The overhead pressure is 0.7 MPa, and the overhead steam is cooled and then flashed to remove the acid gas. At the ammonia-rich point in the middle of the column, the side-drawn steam is fed into a three-stage flash tank, and the side-drawn rate is 14%. The three-stage flash tank includes: a first-stage flash tank with a flash pressure of 0.25 MPa and a flash temperature of 140°C; a second-stage flash tank with a flash pressure of 0.2 MPa and a flash temperature of 110°C; and a third-stage flash tank with a flash pressure of 0.15 MPa and a flash temperature of 50°C, after which the ammonia gas is flashed out by cooling and depressurizing, and the liquid phase from the flash tanks is combined and then returned to the wastewater tank. The deacidification and deamination wastewater from the bottom of the deacidification and deamination column is heated and then fed into the dephenol extraction column.

[0078] The n-butyl ether from the solvent circulation tank is countercurrently extracted with the deacidification and deamination wastewater from the bottom of the deacidification and deamination column to remove phenol, with an extraction temperature of 90°C, an extraction stage (theoretical plate number) of 3, and a volume ratio of n-butyl ether to deacidification and deamination wastewater of 1:4, to obtain a phenol-containing extraction phase and an oil-phenol-ammonia-removed wastewater, which is fed into a subsequent biochemical treatment unit.

[0079] Solvent stripping: the extraction phases from the deoiling extraction column and the dephenol extraction column are combined and then heated by heat exchange with the overhead discharge from the solvent stripping column before being fed into the solvent stripping column, which uses rectification to recover the n-butyl ether, with an overhead temperature of 170°C, a column bottom temperature of 236°C, an overhead pressure of 0.2 MPa, and a theoretical plate number of 28. The n-butyl ether is distilled out of the overhead, cooled by heat exchange, and then fed into the solvent circulation tank for recycling, and the column bottom obtains crude phenol and tar.

[0080] The dephenol extraction column effluent (i.e., the oil-phenol-ammonia-removed wastewater) has a COD content of 5670 mg / L, an oil content of 88 mg / L, a total phenol content of 555 mg / L, and an ammonia content of 122 mg / L.

[0081] The above examples are the preferred embodiments of the present application, but the embodiments of the present application are not limited to the above examples, and any changes, modifications, substitutions, combinations, simplifications made without departing from the spirit and principles of the present application are equivalent replacement methods and are included in the protection scope of the present application.

Claims

1. A method for removing oil phenol and ammonia from high-concentration phenol-containing coal chemical wastewater, characterized in that Specifically comprising the following steps: (1) oil extraction: n-butyl ether and coal chemical wastewater into the oil extraction tower extraction, obtain oil extraction phase and oil extraction raffinate phase; (2) deacidification and deamination: the oil extraction raffinate phase is divided into cold and hot two ways and enters the deacidification and deamination tower, the cold feed enters the upper part of the packing section from the top of the deacidification and deamination tower, the hot feed enters the first tray below the packing section, and alkali liquor is added to the deacidification and deamination tower; acidic gas is obtained at the top of the tower, crude ammonia gas is obtained in the middle part of the tower, and deacidification and deamination wastewater is obtained at the tower bottom; (3) extraction and dephenolization: n-butyl ether and deacidification and deamination wastewater are extracted in the dephenol extraction tower to obtain phenol-containing extraction phase and dephenolization wastewater; (4) solvent stripping: the oil-containing extraction phase of step (1) and the phenol-containing extraction phase of step (3) enter the solvent stripping tower, n-butyl ether is obtained at the top of the tower and is sent to the solvent circulation tank for recycling after heat exchange and cooling, and crude phenol and tar are obtained at the tower bottom; The volume ratio of n-butyl ether to coal chemical wastewater in step (1) is 1:9-1:3; The deacidification and deamination tower in step (2) is a composite tower combining a packed tower and a plate tower; The temperature of the cold feed into the tower in step (2) is 30-60℃, the temperature of the hot feed into the tower is 130-170℃, and the tower bottom pressure is 0.3-0.7MPa; The addition of alkali liquor in step (2) is specifically adding sodium hydroxide solution from the position above the tower bottom to the position below the side line extraction position in the deacidification and deamination tower; The volume ratio of n-butyl ether to deacidification and deamination wastewater in step (3) is 1:4-1:

1.

2. The method according to claim 1, wherein the method is characterized by: The extraction in steps (1) and (3) is countercurrent extraction.

3. The method according to claim 1, wherein the method is characterized by: The extraction temperature in step (1) is 30-90℃; the extraction stage number of the oil extraction tower is 2-8; and the theoretical tray number is the same as the extraction stage number.

4. The method according to claim 1, wherein the method is characterized by: The crude ammonia gas obtained in the middle part of the tower in step (2) is specifically crude ammonia gas extracted from the ammonia gas enrichment zone side line, then enters the three-stage fractionation, and ammonia gas is flashed out after cooling and pressure reduction, and the side line extraction rate is 8%-15%; The three-stage fractionation includes a first-stage flash tank, the flash pressure is 0.25-0.45MPa, and the flash temperature is 120-140℃; a second-stage flash tank, the flash pressure is 0.2-0.4MPa, and the flash temperature is 70-110℃; and a third-stage flash tank, the flash pressure is 0.10-0.25MPa, and the flash temperature is 30-50℃.

5. The method according to claim 1, wherein the method is characterized by: The extraction temperature in step (3) is 30-90℃, and the extraction stage number of the dephenol extraction tower is 3-10; and the theoretical tray number is the same as the extraction stage number.

6. The method according to claim 1, wherein the method is characterized by: The tower top temperature of the solvent stripping tower in step (4) is 125-170℃; the tower bottom temperature is 195-240℃; the tower top pressure is 0.1-0.2MPa; and the theoretical tray number is 15-35.

7. The method according to claim 1, wherein the method is characterized by: The coal chemical wastewater has a COD content of 22500-51200mg / L, an oil content of 2140-3130mg / L, a total phenol content of 6800-15100mg / L, and an ammonia content of 5600-9770mg / L.

8. The method according to claim 7, wherein the method is characterized by: The coal chemical wastewater has a COD content of 24500-48600mg / L, an oil content of 2430-3000mg / L, a total phenol content of 6800-14400mg / L, and an ammonia content of 5600-8800mg / L.

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