CN107890748B - Medium-temperature acidic gas pre-concentration process - Google Patents

Medium-temperature acidic gas pre-concentration process Download PDF

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Publication number
CN107890748B
CN107890748B CN201711029997.3A CN201711029997A CN107890748B CN 107890748 B CN107890748 B CN 107890748B CN 201711029997 A CN201711029997 A CN 201711029997A CN 107890748 B CN107890748 B CN 107890748B
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tower
gas
flash
sulfur
stripping
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CN107890748A (en
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王同宝
周兴
傅亮
严东
褚永良
刘芹
胡有元
陈国平
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Sinopec Engineering Group Co Ltd
Sinopec Ningbo Engineering Co Ltd
Sinopec Ningbo Technology Research Institute
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Sinopec Engineering Group Co Ltd
Sinopec Ningbo Engineering Co Ltd
Sinopec Ningbo Technology Research Institute
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    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
    • B01D53/1425—Regeneration of liquid absorbents
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
    • B01D53/1418—Recovery of products
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/14—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols by absorption
    • B01D53/1456—Removing acid components
    • B01D53/1462—Removing mixtures of hydrogen sulfide and carbon dioxide
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10K—PURIFYING OR MODIFYING THE CHEMICAL COMPOSITION OF COMBUSTIBLE GASES CONTAINING CARBON MONOXIDE
    • C10K1/00—Purifying combustible gases containing carbon monoxide
    • C10K1/002—Removal of contaminants
    • C10K1/003—Removal of contaminants of acid contaminants, e.g. acid gas removal
    • C10K1/004—Sulfur containing contaminants, e.g. hydrogen sulfide
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10K—PURIFYING OR MODIFYING THE CHEMICAL COMPOSITION OF COMBUSTIBLE GASES CONTAINING CARBON MONOXIDE
    • C10K1/00—Purifying combustible gases containing carbon monoxide
    • C10K1/002—Removal of contaminants
    • C10K1/003—Removal of contaminants of acid contaminants, e.g. acid gas removal
    • C10K1/005—Carbon dioxide
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10K—PURIFYING OR MODIFYING THE CHEMICAL COMPOSITION OF COMBUSTIBLE GASES CONTAINING CARBON MONOXIDE
    • C10K1/00—Purifying combustible gases containing carbon monoxide
    • C10K1/08—Purifying combustible gases containing carbon monoxide by washing with liquids; Reviving the used wash liquors

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  • Chemical & Material Sciences (AREA)
  • Engineering & Computer Science (AREA)
  • Combustion & Propulsion (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Organic Chemistry (AREA)
  • Analytical Chemistry (AREA)
  • Gas Separation By Absorption (AREA)
  • Treating Waste Gases (AREA)

Abstract

The invention relates to a medium-temperature acid gas pre-concentration process, which comprises the steps of effectively exchanging heat with a lean absorbent from a thermal regeneration tower before the sulfur-rich absorbent is thermally regenerated, heating the sulfur-rich absorbent to 70-110 ℃ after heat exchange, and then performing flash evaporation operation and gas stripping operation to ensure that CO in the sulfur-rich absorbent can be subjected to CO pre-concentration2Further desorbing the gas to obviously increase the H in the gas2The concentration of S realizes the operation of acid gas concentration under the condition of medium temperature; the sulfur-rich absorbent before entering concentration is subjected to effective heat exchange with the lean absorbent from the thermal regeneration tower through the heat exchanger, so that the effective utilization of energy is realized, and the production cost is reduced; the process is also provided with bypass regulation, so that the temperature of the sulfur-rich absorbent entering the flash tower can be effectively controlled, the sulfur-rich absorbent is ensured to be concentrated at different degrees at proper temperature, and further Claus gases with different concentrations are generated after thermal regeneration to meet the requirements of a downstream sulfur recovery working section.

Description

Medium-temperature acidic gas pre-concentration process
Technical Field
The invention relates to the field of synthesis gas purification, in particular to a medium-temperature acidic gas pre-concentration process.
Background
The synthesis gas purification is mainly used for removing H in the crude synthesis gas2S、COS、CO2Thereby producing a purified gas satisfying the process requirements. The synthesis gas purification processes commonly used in industry are generally classified into a hot process and a cold process according to the difference of absorption temperature. The hot process is most well known as the Selexol process and the MDEA process, while the cold process is represented by the low-temperature methanol washing process. The low-temperature methanol washing process is a synthesis gas purification process jointly developed by German Linde company and Lurgi company in the early 50 s. The process takes methanol as an absorption solvent, utilizes the excellent characteristic that the methanol has extremely high solubility to acid gas at low temperature, and adopts physical absorption to remove the acid gas in raw material gas.
The prior method also lacks the effect of H in the export acid gas2Effective regulation and control means of S concentration. Different sulfur recovery processes for H in acid gases during production operations2The concentration of S has different requirements. When H is contained in the acid gas2When the concentration of S is low, the sulfur recovery device cannot normally operate for a long time. The prior art has been to increase the H in the acid gas by recycling a portion of the acid gas back to the upstream system2S concentration, however, for a partially constructed syngas purification plant, increasing H in the acid gas by increasing the acid gas circulation, is limited by process settings, inlet syngas sulfur content fluctuations, and the like2The extra investment and energy consumption required for the concentration of S are higher, and the economy is poor.
Disclosure of Invention
The invention aims to solve the technical problem of providing a method which has low energy consumption and can obviously improve H in acid gas at the outlet of a system aiming at the current situation of the prior art2S concentration, facilitating H regulation2A medium-temperature acid gas pre-concentration process with S concentration.
The technical scheme adopted by the invention for solving the technical problems is as follows: a medium-temperature acid gas pre-concentration process comprises the following steps:
rich sulfur uptake from upstreamThe temperature of the collecting agent is-20 to-60 ℃, and the pressure is 0.6 to 1.5MPaG and H2The mol content of S is 0.05-5%, the S and lean absorbent from a thermal regeneration tower are subjected to heat exchange and temperature rise to 70-110 ℃, preferably 90-100 ℃, and then the S and lean absorbent are sent to a flash tower to be subjected to reduced pressure flash evaporation, and the pressure in the flash tower is controlled to be 0.15-0.45 MPaG;
flash evaporated CO-containing2And part H2Conveying the flash evaporation gas of the S from the top of the flash evaporation tower to a downstream process, and conveying the sulfur-rich absorbent at the bottom of the flash evaporation tower to a stripping tower for stripping operation;
introducing nitrogen at the temperature of 10-110 ℃ into the lower part of the stripping tower, and controlling the pressure in the stripping tower to be 0.15-0.45 MPaG; CO in sulfur-rich absorbent in stripping operation2And a small amount of H2S is further desorbed, the gas stripped out is converged with the flash evaporation gas, the cooled gas and the flash evaporation gas are sent to a downstream process, and the sulfur-rich absorbent stripped is discharged from the bottom of a stripping tower;
conveying the sulfur-rich absorbent discharged from the bottom of the stripping tower to a thermal regeneration tower for thermal regeneration, and controlling the operating temperature of the thermal regeneration tower to be 80-105 ℃ and the operating pressure to be 0.20-0.30 MPaG; desorbing the dissolved acid gas sulfide in the rich absorbent to form Claus gas, outputting the Claus gas from the top of the thermal regeneration tower, cooling to 30 deg.C, wherein H in the Claus gas2The molar content of S is 40-80%, the gas is sent to a downstream sulfur recovery process, and the condensate is separated and then returned to the top of the thermal regeneration tower as a reflux;
outputting the lean absorbent after the thermal regeneration from the lower part of the thermal regeneration tower, exchanging heat with the sulfur-rich absorbent, and then sending the lean absorbent to a downstream acid gas absorption process; the absorbent aqueous solution is output from the bottom of the thermal regeneration tower and sent to the downstream.
As an improvement, an inlet of the flash tower is connected with a first liquid inlet pipeline for conveying the sulfur-rich absorbent to be treated, the first liquid inlet pipeline is sequentially connected with a first heat exchanger and a second heat exchanger in series, the flash tower is provided with a first gas phase outlet, the first gas phase outlet is connected with a first gas outlet pipeline, and the first gas outlet pipeline is provided with a control valve for adjusting the flash pressure in the flash tower; the liquid phase outlet of the flash tower is connected with the inlet of the stripping tower; the lower part of the gas stripping tower is connected with a gas stripping gas conveying pipeline, the upper part of the gas stripping tower is provided with a second gas phase outlet, and a liquid phase outlet of the gas stripping tower is connected with an inlet of the thermal regeneration tower. The sulfur-rich absorbent continuously exchanges heat with the lean absorbent from the thermal regeneration tower for two times through the first heat exchanger and the second heat exchanger, and then directly enters the flash tower for flash evaporation operation, so that the effective utilization of energy is realized.
As an improvement, the thermal regeneration tower comprises two parts, namely a section I and a section II, which are arranged up and down, the upper part of the section I of the thermal regeneration tower is communicated with a liquid phase outlet of the stripping tower through a second liquid inlet pipeline, the lower part of the section I of the thermal regeneration tower is connected with a first liquid outlet pipeline for outputting lean absorbent, the lower part of the section II of the thermal regeneration tower is provided with a reboiler for providing heat for regeneration, and the bottom of the thermal regeneration tower is connected with a second liquid outlet pipeline for conveying absorbent aqueous solution.
Preferably, the molar flow ratio of the stripping nitrogen to the sulfur-rich absorbent entering the stripping tower is 1: 10000-450: 10000.
As an improvement, a first bypass is arranged on the first liquid outlet pipe, two ports of the first bypass are respectively positioned at the upstream and the downstream of the second heat exchanger, and a first valve is arranged on the first bypass; and a second bypass is arranged on the sulfur-rich absorbent conveying pipeline, two ports of the second bypass are respectively positioned at the upstream and the downstream of the first heat exchanger, a second valve is arranged on the second bypass, and the first valve and the second valve are respectively used for controlling the flow rate of the lean absorbent entering the first heat exchanger and the flow rate of the sulfur-rich absorbent entering the second heat exchanger, so as to control the flash evaporation temperature of the sulfur-rich absorbent entering the flash tower. Through setting up above-mentioned bypass and adjusting, can control the temperature that gets into the rich sulphur absorbent of flash column effectively, carry out flash distillation concentration of different degrees, and then make it produce the requirement that the different concentration's claus gas satisfies the downstream sulphur recovery workshop section after thermal regeneration.
Preferably, the flow ratio of the first bypass to the main route where the first bypass is located is 0-1: 3; the flow ratio of the second bypass to the main path where the second bypass is located is 0-1: 3. The flow of the first bypass and the second bypass can be controlled according to the actual production condition, and the first valve of the first bypass and the second valve of the second bypass can be closed when necessary, so that the maximum heat exchange is achieved.
As an improvement, the flash tower and the stripper share the same tower body, the flash tower and the stripper are isolated by a first seal head, the flash tower is located above the stripper, a liquid phase outlet of the flash tower is connected with an inlet of the stripper through an external pipeline, a liquid phase outlet of the stripper is connected with an inlet of the thermal regeneration tower through a first circulating pump, the operating temperature in the flash tower is 90-100 ℃, the flash pressure is 0.30-0.40 MPaG, the operating temperature in the stripper is 90-100 ℃, and the operating pressure is 0.30-0.40 MPaG. The sulfur-rich absorbent flows into a stripping tower through an external pipeline under the action of self gravity and steam pressure in the flash tower after flash evaporation treatment, and a liquid phase is pressurized by a first circulating pump and then is introduced into a thermal regeneration tower after steam stripping treatment, so that the fluid conveying cost is saved, and the energy is saved; meanwhile, the two devices are stacked together, so that the occupied area is small, and on the other hand, the flash evaporation operation and the steam stripping operation can be carried out step by step; of course, when the operating pressure in the flash column is sufficiently high, the first circulation pump may also be omitted and the liquid phase may be sent directly to the hot regenerator inlet.
Or the flash tower and the stripping tower share the same tower body, the flash tower and the stripping tower are separated by adopting a tower tray, and the flash tower is positioned above the stripping tower;
the tower tray is provided with a riser for the gas phase in the gas stripping tower to enter the flash tower, the riser forms a second gas phase outlet at the upper part of the gas stripping tower, the tower tray is also provided with a downcomer for the liquid phase in the flash tower to enter the gas stripping tower, the downcomer is communicated with the liquid phase outlet of the flash tower and the inlet of the gas stripping tower, the liquid phase outlet of the gas stripping tower is connected with the inlet of the heat regeneration tower through a first circulating pump, wherein the operating temperature in the flash tower is 90-100 ℃, the flash pressure is 0.30-0.40 MPaG, the operating temperature in the gas stripping tower is 90-100 ℃, and the operating pressure is 0.30-0.40 MPaG. The flash tower and the stripping tower adopt the same tower body structure, and are separated by a tower tray provided with a riser and a downcomer, so that the fluid conveying cost is saved, the energy is saved, and the flash operation and the steam stripping operation in the flash tower can be simultaneously carried out.
Or the flash tower, the stripping tower and the thermal regeneration tower share the same tower body, the flash tower and the stripping tower are isolated by a first end enclosure, the stripping tower and the thermal regeneration tower are isolated by a second end enclosure, and the stripping tower and the thermal regeneration tower of the flash tower are sequentially arranged from top to bottom; the liquid phase outlet of the flash tower is connected with the inlet of the stripping tower through an external pipeline; wherein the operating temperature in the flash tower is 90-100 ℃, the flash pressure is 0.30-0.40 MPaG, the operating temperature in the gas stripping tower is 90-100 ℃, the operating pressure is 0.30-0.40 MPaG, the operating temperature in the thermal regeneration tower is 87-104 ℃, and the operating pressure is 0.22-0.28 MPaG. The sulfur-rich absorbent flows into the stripping tower through an external pipeline under the action of self gravity and steam pressure in the flash tower after flash evaporation operation, and liquid phase is led into the thermal regeneration tower below after the stripping operation, so that the first circulating pump is omitted, and the production cost is saved.
Or the flash tower, the stripping tower and the thermal regeneration tower share the same tower body; the flash tower, the stripping tower and the thermal regeneration tower are sequentially arranged from top to bottom;
the flash tower is separated from the stripping tower through a tower tray, and the stripping tower is separated from the thermal regeneration tower through a second seal head; the tower tray is provided with a riser for the gas phase in the gas stripping tower to enter the flash tower, the riser forms a second gas phase outlet at the upper part of the gas stripping tower, the tower tray is also provided with a downcomer for the liquid phase in the flash tower to enter the gas stripping tower, and the downcomer is communicated with the liquid phase outlet of the flash tower and the inlet of the gas stripping tower, wherein the operating temperature in the flash tower is 90-100 ℃, the flash pressure is 0.30-0.40 MPaG, the operating temperature in the gas stripping tower is 90-100 ℃, and the operating pressure is 0.30-0.40 MPaG. The flash tower, the gas stripping tower and the thermal regeneration tower share the same tower body structure, a fluid pressurizing and conveying device of a first circulating pump between the gas stripping tower and the thermal regeneration tower can be omitted, the fluid conveying cost is reduced to a certain degree, the purposes of energy saving and consumption reduction are achieved, and meanwhile, the three devices are stacked together, so that the occupied area is greatly saved.
Compared with the prior art, the invention has the advantages that: the method comprises the steps of performing effective heat exchange with a lean absorbent from a heat regeneration tower before the sulfur-rich absorbent performs heat regeneration, heating the sulfur-rich absorbent to 70-110 ℃ after heat exchange, and then performing flash evaporation operation and gas stripping operation to enable CO in the sulfur-rich absorbent to be in a CO state2Further desorbing the gas to obviously increase the H in the gas2The concentration of S realizes the operation of acid gas concentration under the condition of medium temperature; on the other hand, the sulfur-rich absorbent before entering concentration is subjected to effective heat exchange with the lean absorbent from the thermal regeneration tower through the first heat exchanger and the second heat exchanger, so that the effective utilization of energy is realized, and the production cost is reduced; the first bypass is arranged on the first liquid inlet pipeline, the second bypass is arranged on the second liquid inlet pipeline, the first valve is arranged on the first bypass, the second valve is arranged on the second bypass, the temperature of the sulfur-rich absorbent entering the flash tower can be effectively controlled, the sulfur-rich absorbent is ensured to be concentrated at different degrees at proper temperature, and then Claus gases with different concentrations generated after thermal regeneration meet the requirements of a downstream sulfur recovery working section; and finally, the flash tower, the stripping tower and the heat regeneration tower share the structure of the same tower body, so that on one hand, the occupied area is greatly reduced, on the other hand, a first circulating pump device can be omitted, the fluid conveying cost is reduced to a certain degree, and the purposes of energy conservation and consumption reduction are achieved.
The invention is particularly suitable for being used in a matched gas purification process and is used for regenerating the absorbent in the gas purification process.
Drawings
FIG. 1 is a schematic view of example 1 of the present invention;
FIG. 2 is a schematic view of example 2 of the present invention;
FIG. 3 is a schematic view of example 3 of the present invention;
fig. 4 is a schematic view of embodiment 4 of the present invention.
Detailed Description
The invention is described in further detail below with reference to the accompanying examples.
The gases treated in the following examples are all CO-rich gases from an upstream gas production plant2And H2Non-purified gas of acid gas such as S, etc. with main gas composition of H275-8 percent of CO, 75-0.1 percent of CO2Content 60% -2%, H2The S content is 5-0.0001%, and the used absorbent can be Selexol, NHD, MDEA, methanol and other physical absorbents.
Examples 1,
In the medium-temperature acid gas pre-concentration process shown in FIG. 1, the sulfur-rich absorbent from upstream has a flow rate of 274954kg/H, a temperature of-55 ℃ and a pressure of 1.28MPaG and H2The mol content of S is 0.318%, and the S exchanges heat with lean absorbent from the thermal regeneration tower 1 sequentially through a first heat exchanger 5 and a second heat exchanger 4 to be heated to 95 ℃, and then is sent to a flash tower 2 to be subjected to reduced pressure flash evaporation; specifically, the first heat exchanger 5 and the second heat exchanger 4 may be connected in series, the sulfur-rich absorbent passes through the tube side of the first heat exchanger 5 and the tube side of the second heat exchanger 4, and correspondingly, the lean absorbent passes through the shell side of the first heat exchanger 5 and the shell side of the second heat exchanger 4; of course, the sulfur-rich absorbent can also reach the same purpose of heat exchange by taking the sulfur-rich absorbent away from the shell side of the first heat exchanger 5 and the second heat exchanger 4, taking the lean absorbent away from the tube side of the second heat exchanger 4 and the first heat exchanger 5, or respectively taking the sulfur-rich absorbent and the lean absorbent across the tube side and the shell side of the first heat exchanger 5 and the second heat exchanger 4.
A first gas phase outlet is arranged at the top of the flash tower 2, the first gas phase outlet is connected with a first gas outlet pipeline, a control valve 14 for adjusting the flash pressure in the flash tower 2 is arranged on the first gas outlet pipeline, and the flash pressure in the embodiment is 0.35 MPaG; flash evaporated CO-containing2And part H2The flash evaporation gas of S is sent to a downstream process from the top of the flash evaporation tower 2 through a first gas outlet pipeline, and the sulfur-rich absorbent after flash evaporation flows out from the bottom of the flash evaporation tower 2 and is sent to a stripper 3 for stripping operationDo this.
In this embodiment, the same tower body is shared by the flash tower 2 and the stripper tower 3, the first seal head 21 is adopted for isolation between the two, the flash tower 2 is positioned above the stripper tower 3, the liquid phase outlet of the flash tower 2 is connected with the inlet of the stripper tower 3 through an external pipeline, the sulfur-rich absorbent can flow into the stripper tower 3 through the external pipeline under the action of self gravity and the steam pressure in the flash tower 2 after flash evaporation treatment, thus, the fluid conveying cost is saved, the energy is saved, meanwhile, the two devices are stacked together, and the occupied area is also saved.
The lower part of the stripping tower 3 is connected with a stripping gas conveying pipeline, nitrogen enters the stripping tower 3 through the stripping gas conveying pipeline, and the flow of the stripping nitrogen in the embodiment is 310Nm3The temperature is 40 ℃, the pressure of stripping nitrogen is 0.40MPaG, and CO in the sulfur-rich absorbent is in the stripping operation2And a small amount of H2And S is further desorbed, a second gas phase outlet is arranged at the upper part of the gas stripping tower 3, the gas-stripped mixed gas is discharged through the second gas phase outlet, is converged with the flash steam and then is conveyed to the downstream, and the gas-stripped sulfur-rich absorbent is output from the bottom of the gas stripping tower 3 and is conveyed to the thermal regeneration tower 1.
The temperature of the sulfur-rich absorbent output from the bottom of the stripper 3 is 85 ℃, the sulfur-rich absorbent is conveyed to the upper part of the thermal regeneration tower 1 through a first circulating pump 11, the thermal regeneration tower 1 comprises an I section and an II section which are arranged up and down, the two sections are separated through a lift cap, when regeneration is carried out, the operating temperature of the thermal regeneration tower 1 is 87-104 ℃, the operating pressure is 0.22-0.28 MPaG, a reboiler 8 connected with the lower part of the II section of the thermal regeneration tower 1 can provide heat for thermal regeneration of the thermal regeneration tower 1 through an indirect heat exchange mode, acid gas dissolved in the rich absorbent is desorbed to form Claus gas, the Claus gas is output from the top of the thermal regeneration tower 1 and cooled to the temperature of 30 ℃, wherein the H in the Claus gas2The molar content of S is 40%, the cooled claus gas is sent to the downstream sulfur recovery process, and the condensate is separated and returned to the top of the thermal regeneration column 1 as reflux.
The lower part of the section I of the thermal regeneration tower 1 is connected with a first liquid outlet pipeline, the first liquid outlet pipeline outputs the thermally regenerated lean absorbent through a second circulating pump 10, the temperature of the lean absorbent is 102 ℃, the pressure of the lean absorbent is 0.28MPaG, and then the lean absorbent is subjected to heat exchange with the sulfur-rich absorbent sequentially through the second heat exchanger 4 and the first heat exchanger 5 and then sent to an upstream acid gas absorption process; the bottom of the section II of the thermal regeneration tower 1 is connected with a second liquid outlet pipeline, and the absorbent aqueous solution is output from the second liquid outlet pipeline and sent to the purification and separation process of the downstream absorbent.
A first bypass is arranged on the first liquid outlet pipe, two ports of the first bypass are respectively positioned at the upstream and the downstream of the second heat exchanger 4, and a first valve 12 is arranged on the first bypass; the sulfur-rich absorbent conveying pipeline is provided with a second bypass, two ports of the second bypass are respectively located at the upstream and the downstream of the first heat exchanger 5, the second bypass is provided with a second valve 13, the first valve 12 and the second valve 13 are respectively used for controlling the flow rate of the lean absorbent entering the first heat exchanger 5 and the flow rate of the sulfur-rich absorbent entering the second heat exchanger 4, and further controlling the flash evaporation temperature of the sulfur-rich absorbent entering the flash tower 2 when the load of the device changes, in the embodiment, the flow rate of the first bypass under the normal load is 0, and the flow rate of the second bypass under the normal load is 0.
When H is contained in the acid gas2When the concentration of S is low, the prior art adopts a mode of recycling part of acid gas back to an upstream system to increase H in the acid gas2(ii) the concentration of S; in the acid gas satisfying the same export H2On the basis of the S concentration, compared with the process of pressurizing and circulating the acid gas to the upstream, the method for concentrating the acid gas has the advantages of less investment, lower energy consumption and lower operation cost. Taking the same-scale low-temperature methanol washing device in the prior art as an example, when H in the acid gas is discharged2When the concentration of S is kept at 40%, the consumption, investment and operation cost of the utility project by adopting the acid gas pressurization circulation flow and the process flow in the embodiment 1 are compared as follows:
contrast item Example 1 comparison with existing concentration techniques
Power consumption (Kw) -140
Circulating water consumption (t/h) -20
0.4MPag steam consumption (t/h) -0.4
Consumption of Nitrogen (Nm)3/h) +310
Investment of equipment (Wanyuan) -30
Annual operating expenses (Wanyuan) -55
Note: 1. "-" indicates decrease and "+" indicates increase;
2. unit cost of utility consumption: electricity 0.6 yuan/kw; circulating water is 0.3 yuan/ton; steam is 100 yuan/ton; nitrogen 0.2 yuan/Nm3。
Examples 2,
As shown in FIG. 2, the medium temperature acid gas pre-concentration process is provided, wherein the sulfur-rich absorbent from upstream has a flow rate of 274954kg/H, a temperature of-55 ℃ and a pressure of 1.28MPaG and H2The molar content of S is 0.318%;
the difference in this embodiment from embodiment 1 is that the flash column 2 and the stripper column 3 share the same column body, the two are separated by a tray 9, and the flash column 2 is located above the stripper column 3; a gas riser 91 for allowing the gas phase in the stripping tower 3 to enter the flash tower 2 is arranged on the tray 9, the gas riser 91 forms a second gas phase outlet at the upper part of the stripping tower 3 in the embodiment 1, a downcomer 92 for allowing the liquid phase in the flash tower 2 to enter the stripping tower 3 is also arranged on the tray 9, the downcomer 92 is communicated with the liquid phase outlet of the flash tower 2 and the inlet of the stripping tower 3, and the liquid phase outlet of the stripping tower 3 is connected with the inlet of the thermal regeneration tower 1 through a first circulating pump 11; the rest of the present embodiment is the same as embodiment 1, and will not be described in detail, and the corresponding process operation parameters thereof are changed as follows:
the temperature of the sulfur-rich absorbent entering the flash tower 2 after heat exchange through the first heat exchanger 5 and the second heat exchanger 4 is 95 ℃;
the flash pressure of the flash column 2 was 0.35 MPaG;
the flow rate of stripping nitrogen gas is 310Nm3The temperature of stripping nitrogen is 40 ℃, and the pressure of the stripping nitrogen is 0.4 MPaG;
the temperature of the sulfur-rich absorbent output from the bottom of the stripper 3 is 85 ℃;
the operation temperature of the thermal regeneration tower 1 is 87-104 ℃, and the operation pressure is 0.22-0.28 MPaG;
the temperature of the claus gas output from the top of the thermal regeneration tower 1 after being cooled by the cooler is 30 ℃.
H in Claus gas2The molar content of S is 40 percent;
taking the same-scale low-temperature methanol washing device in the prior art as an example, when H in the acid gas is discharged2When the concentration of S is kept at 40%, the consumption, investment and operation cost of the public works of the acid gas pressurization circulation flow adopted in the prior art and the process flow adopted in the embodiment are compared as follows:
contrast item Example 2 comparison with the existing concentration technology
Power consumption (Kw) -140
Circulating water consumption (t/h) -20
0.4MPag steam consumption (t/h) -0.4
Consumption of Nitrogen (Nm)3/h) +310
Investment of equipment (Wanyuan) -30
Annual operating expenses (Wanyuan) -55
Note: 1. "-" indicates decrease and "+" indicates increase;
2. unit cost of utility consumption: electricity 0.6 yuan/kw; circulating water is 0.3 yuan/ton; steam is 100 yuan/ton; nitrogen 0.2 yuan/Nm3。
Examples 3,
As shown in FIG. 3, the medium temperature acid gas pre-concentration process is provided, wherein the sulfur-rich absorbent from upstream has a flow rate of 274954kg/H, a temperature of-55 ℃ and a pressure of 1.28MPaG and H2The molar content of S is 0.318%;
the difference in this embodiment from embodiment 1 is that the flash tower 2, the stripper tower 3 and the thermal regeneration tower 1 share the same tower body, the flash tower 2 is isolated from the stripper tower 3 by the first head 21, the stripper tower 3 is isolated from the thermal regeneration tower 1 by the second head 31, and the flash tower 2, the stripper tower 3 and the thermal regeneration tower 1 are sequentially arranged from top to bottom; the liquid phase outlet of the flash tower 2 is connected with the inlet of the gas stripping tower 3 through an external fluid pipeline, the liquid phase outlet of the gas stripping tower 3 is connected with the inlet of the heat regeneration tower 1 through a second liquid inlet pipeline, under the action of self gravity and the internal pressure of the tower body, the sulfur-rich absorbent enters the gas stripping tower 3 from the flash tower 2 and then enters the second liquid inlet pipeline, exchanges heat through the second heat exchanger 4 and then is guided into the heat regeneration tower 1 for heat regeneration treatment. Thus, a fluid pressurizing and conveying device of the first circulating pump 11 between the stripping tower 3 and the heat regeneration unit is omitted, and the production cost is reduced. Otherwise, the rest of the present embodiment is the same as embodiment 1, and is not detailed again, and the corresponding process operation parameters are changed as follows:
the temperature of the sulfur-rich absorbent entering the flash tower 2 after heat exchange through the first heat exchanger 5 and the second heat exchanger 4 is 95 ℃;
the flash pressure of the flash column 2 was 0.35 MPaG;
the flow rate of stripping nitrogen gas is 310Nm3The temperature of stripping nitrogen is 40 ℃, and the pressure of the stripping nitrogen is 0.4 MPaG;
the temperature of the sulfur-rich absorbent output from the bottom of the stripper 3 is 85 ℃;
the operation temperature of the thermal regeneration tower 1 is 87-104 ℃, and the operation pressure is 0.22-0.28 MPaG;
the temperature of the Claus gas output from the top of the thermal regeneration tower 1 after being cooled by a cooler is 30 ℃;
h in Claus gas2The molar content of S is 40%.
Taking the same-scale low-temperature methanol washing device in the prior art as an example, when H in the acid gas is discharged2When the concentration of S is kept at 40%, the consumption, investment and operation cost of the public works of the acid gas pressurization circulation flow adopted in the prior art and the process flow adopted in the embodiment are compared as follows:
contrast item Example 3 comparison with existing concentration techniques
Power consumption (Kw) -140
Circulating water consumption (t/h) -20
0.4MPag steam consumption (t/h) -0.4
Consumption of Nitrogen (Nm)3/h) +310
Investment of equipment (Wanyuan) -30
Annual operating expenses (Wanyuan) -55
Note: 1. "-" indicates decrease and "+" indicates increase;
2. unit cost of utility consumption: electricity 0.6 yuan/kw; circulating water is 0.3 yuan/ton; steam is 100 yuan/ton; nitrogen 0.2 yuan/Nm3。
Examples 4,
As shown in FIG. 4, the medium temperature acid gas pre-concentration process is provided, wherein the sulfur-rich absorbent from upstream has a flow rate of 274954kg/H, a temperature of-55 ℃ and a pressure of 1.28MPaG and H2The molar content of S is 0.318%;
the difference in this embodiment from embodiment 1 is that the flash column 2, the stripper column 3 and the thermal regeneration column 1 share the same column body, the flash column 2 is isolated from the stripper column 3 by the tray 9, the stripper column 3 is isolated from the thermal regeneration column 1 by the second head 31, and the flash column 2, the stripper column 3 and the thermal regeneration column 1 are sequentially arranged from top to bottom; be equipped with gas riser 91 and downcomer 92 on tray 9, gas in the stripper 3 directly enters into the flash column 2 of top through gas riser 91 in, correspondingly, the liquid phase in the flash column 2 enters into the stripper 3 of below through downcomer 92 in, can avoid like this to increase the pipeline of connecting flash column 2 and stripper 3 outside the tower body, also makes flash process and gas stripping process go on simultaneously. The liquid phase outlet of the gas stripping tower 3 is connected with the inlet of the heat regeneration tower 1 through a second liquid inlet pipeline, under the action of self gravity and the internal pressure of the tower body, the sulfur-rich absorbent enters the gas stripping tower 3 from the flash tower 2 through the downcomer 92, then enters the second liquid inlet pipeline, exchanges heat through the second heat exchanger 4, and then is guided into the heat regeneration tower 1 for heat regeneration treatment. Thus, a fluid pressurizing and conveying device of the first circulating pump 11 between the stripping tower 3 and the heat regeneration unit is omitted, and the production cost is reduced. Otherwise, the rest of the present embodiment is the same as embodiment 1, and is not detailed again, and the corresponding process operation parameters are changed as follows:
the temperature of the sulfur-rich absorbent entering the flash tower 2 after heat exchange through the first heat exchanger 5 and the second heat exchanger 4 is 95 ℃;
the flash pressure of the flash column 2 was 0.35 MPaG;
the flow rate of stripping nitrogen gas is 310Nm3The temperature of stripping nitrogen is 40 ℃, and the pressure of the stripping nitrogen is 0.4 MPaG;
the temperature of the sulfur-rich absorbent output from the bottom of the stripper 3 is 85 ℃;
the operation temperature of the thermal regeneration tower 1 is 87-104 ℃, and the operation pressure is 0.22-0.28 MPaG;
the temperature of the Claus gas output from the top of the thermal regeneration tower 1 after being cooled by a cooler is 30 ℃;
h in Claus gas2The molar content of S is 40%.
Taking the same-scale low-temperature methanol washing device in the prior art as an example, when H in the acid gas is discharged2When the concentration of S is kept at 40%, the consumption, investment and operation cost of the public works of the acid gas pressurization circulation flow adopted in the prior art and the process flow adopted in the embodiment are compared as follows:
contrast item Example 4 comparison with existing concentration techniques
Power consumption (Kw) -140
Circulating water consumption (t/h) -20
0.4MPag steam consumption (t/h) -0.4
Consumption of Nitrogen (Nm)3/h) +310
Investment of equipment (Wanyuan) -30
Annual operating expenses (Wanyuan) -55
Note: 1. "-" indicates decrease and "+" indicates increase;
2. unit cost of utility consumption: electricity 0.6 yuan/kw; circulating water is 0.3 yuan/ton; steam is 100 yuan/ton; nitrogen 0.2 yuan/Nm3。
In summary, when H is in the acid gas2When the concentration of S is low, the method is similar to the method of recycling part of the acid gas to the upstream system in the prior art to increase the H in the acid gas2Comparing the operation modes of S concentration; in the acid gas satisfying the same export H2On the basis of the concentration of S, the device investment, the production operation cost and the energy consumption of the acidic gas concentration method are greatly reduced.

Claims (2)

1. A medium-temperature acid gas pre-concentration process is characterized by comprising the following steps:
the temperature of the sulfur-rich absorbent from upstream is-20 to-60 ℃, and the pressure is 0.6 to 1.5MPaG and H2The molar content of S is 0.05-5%, and the S and lean absorbent from the thermal regeneration tower (1) are subjected to heat exchange and temperature rise to 70-110 ℃, and then are sent to the flash tower (2) for reduced pressure flash evaporation;
flash evaporated CO-containing2And part H2Conveying the flash evaporation gas of the S from the top of the flash evaporation tower (2) to a downstream process, and conveying the sulfur-rich absorbent at the bottom of the flash evaporation tower (2) to a stripping tower (3) for stripping operation;
introducing nitrogen at the temperature of 10-110 ℃ into the lower part of the stripping tower (3); CO in sulfur-rich absorbent in stripping operation2And a small amount of H2S is further desorbed, the gas stripped out is converged with the flash evaporation gas, the cooled gas and the flash evaporation gas are sent to a downstream process, and the sulfur-rich absorbent stripped is discharged from the bottom of a stripping tower (3);
the sulfur-rich absorbent discharged from the bottom of the stripping tower (3) is sent to a thermal regeneration tower (1) for thermal regeneration, the operating temperature of the thermal regeneration tower (1) is controlled to be 80-105 ℃, and the operating pressure is controlled to be 0.20-0.30 MPaG; desorbing the dissolved acid gas sulfide in the rich absorbent to form Claus gas, outputting the Claus gas from the top of the thermal regeneration tower (1), and cooling to the temperature of 30 ℃, wherein H in the Claus gas2The molar content of S is 40-80%, the gas is sent to a downstream sulfur recovery process, and the condensate is separated and then returned to the top of the thermal regeneration tower (1) as a reflux;
the lean absorbent after the heat regeneration is output from the lower part of the heat regeneration tower (1), exchanges heat with the sulfur-rich absorbent and then is sent to a downstream acid gas absorption process; the absorbent aqueous solution is output from the bottom of the thermal regeneration tower (1) and sent to the downstream;
an inlet of the flash tower (2) is connected with a first liquid inlet pipeline for conveying the sulfur-rich absorbent to be treated, a first heat exchanger (5) and a second heat exchanger (4) are sequentially connected in series on the first liquid inlet pipeline, a first gas phase outlet is arranged on the flash tower (2), the first gas phase outlet is connected with a first gas outlet pipeline, and a control valve (14) for adjusting the flash pressure in the flash tower (2) is arranged on the first gas outlet pipeline; the liquid phase outlet of the flash tower (2) is connected with the inlet of the stripping tower (3); the lower part of the stripping tower (3) is connected with a stripping gas conveying pipeline, the upper part of the stripping tower (3) is provided with a second gas phase outlet, and a liquid phase outlet of the stripping tower (3) is connected with an inlet of the thermal regeneration tower (1);
the heat regeneration tower (1) comprises a section I and a section II which are arranged up and down, the upper part of the section I of the heat regeneration tower (1) is communicated with a liquid phase outlet of the stripping tower (3) through a second liquid inlet pipeline, the lower part of the section I of the heat regeneration tower (1) is connected with a first liquid outlet pipeline for outputting a lean absorbent, the lower part of the section II of the heat regeneration tower (1) is provided with a reboiler (8) for providing heat regeneration heat, and the bottom of the heat regeneration tower (1) is connected with a second liquid outlet pipeline for conveying an absorbent aqueous solution;
the ratio of the molar flow of the stripping nitrogen to the molar flow of the sulfur-rich absorbent entering the stripping tower (3) is (1: 10000) - (450: 10000);
a first bypass is arranged on the first liquid outlet pipe, two ports of the first bypass are respectively positioned at the upstream and the downstream of the second heat exchanger (4), and a first valve (13) is arranged on the first bypass; a second bypass is arranged on the sulfur-rich absorbent conveying pipeline, two ports of the second bypass are respectively positioned at the upstream and the downstream of the first heat exchanger (5), a second valve (12) is arranged on the second bypass, and the first valve (13) and the second valve (12) are respectively used for controlling the flow rate of the lean absorbent entering the second heat exchanger (4) and the flow rate of the sulfur-rich absorbent entering the first heat exchanger (5), so that the flash evaporation temperature of the sulfur-rich absorbent entering the flash tower (2) is controlled;
the flow ratio of the first bypass to the main road where the first bypass is located is 0-1: 3; the flow ratio of the second bypass to a main road where the second bypass is located is 0-1: 3;
the heat regeneration tower is characterized in that the flash tower (2) and the stripping tower (3) share the same tower body, the flash tower (2) and the stripping tower (3) are isolated by a first seal head (21), the flash tower (2) is located above the stripping tower (3), a liquid phase outlet of the flash tower (2) is connected with an inlet of the stripping tower (3) through an external pipeline, a liquid phase outlet of the stripping tower (3) is connected with an inlet of the heat regeneration tower (1) through a first circulating pump (11), the operating temperature in the flash tower (2) is 90-100 ℃, the flash pressure is 0.30-0.40 MPaG, the operating temperature in the stripping tower (3) is 90-100 ℃, and the operating pressure is 0.30-0.40 MPaG.
2. A medium-temperature acid gas pre-concentration process is characterized by comprising the following steps:
the temperature of the sulfur-rich absorbent from upstream is-20 to-60 ℃, and the pressure is 0.6 to 1.5MPaG and H2The molar content of S is 0.05-5%, and the S and lean absorbent from the thermal regeneration tower (1) are subjected to heat exchange and temperature rise to 70-110 ℃, and then are sent to the flash tower (2) for reduced pressure flash evaporation;
flash evaporated CO-containing2And part H2Conveying the flash evaporation gas of the S from the top of the flash evaporation tower (2) to a downstream process, and conveying the sulfur-rich absorbent at the bottom of the flash evaporation tower (2) to a stripping tower (3) for stripping operation;
introducing nitrogen at the temperature of 10-110 ℃ into the lower part of the stripping tower (3); CO in sulfur-rich absorbent in stripping operation2And a small amount of H2S is further desorbed, the gas stripped out is converged with the flash evaporation gas, the cooled gas and the flash evaporation gas are sent to a downstream process, and the sulfur-rich absorbent stripped is discharged from the bottom of a stripping tower (3);
the sulfur-rich absorbent discharged from the bottom of the stripping tower (3) is sent to a thermal regeneration tower (1) for thermal regeneration; desorbing the dissolved acid gas sulfide in the rich absorbent to form Claus gas, outputting the Claus gas from the top of the thermal regeneration tower (1), and cooling to the temperature of 30 ℃, wherein H in the Claus gas2The molar content of S is 40-80%, the gas is sent to a downstream sulfur recovery process, and the condensate is separated and then returned to the top of the thermal regeneration tower (1) as a reflux;
the lean absorbent after the heat regeneration is output from the lower part of the heat regeneration tower (1), exchanges heat with the sulfur-rich absorbent and then is sent to a downstream acid gas absorption process; the absorbent aqueous solution is output from the bottom of the thermal regeneration tower (1) and sent to the downstream;
an inlet of the flash tower (2) is connected with a first liquid inlet pipeline for conveying the sulfur-rich absorbent to be treated, a first heat exchanger (5) and a second heat exchanger (4) are sequentially connected in series on the first liquid inlet pipeline, a first gas phase outlet is arranged on the flash tower (2), the first gas phase outlet is connected with a first gas outlet pipeline, and a control valve (14) for adjusting the flash pressure in the flash tower (2) is arranged on the first gas outlet pipeline; the liquid phase outlet of the flash tower (2) is connected with the inlet of the stripping tower (3); the lower part of the stripping tower (3) is connected with a stripping gas conveying pipeline, the upper part of the stripping tower (3) is provided with a second gas phase outlet, and a liquid phase outlet of the stripping tower (3) is connected with an inlet of the thermal regeneration tower (1);
the heat regeneration tower (1) comprises a section I and a section II which are arranged up and down, the upper part of the section I of the heat regeneration tower (1) is communicated with a liquid phase outlet of the stripping tower (3) through a second liquid inlet pipeline, the lower part of the section I of the heat regeneration tower (1) is connected with a first liquid outlet pipeline for outputting a lean absorbent, the lower part of the section II of the heat regeneration tower (1) is provided with a reboiler (8) for providing heat regeneration heat, and the bottom of the heat regeneration tower (1) is connected with a second liquid outlet pipeline for conveying an absorbent aqueous solution;
the ratio of the molar flow of the stripping nitrogen to the molar flow of the sulfur-rich absorbent entering the stripping tower (3) is (1: 10000) - (450: 10000);
a first bypass is arranged on the first liquid outlet pipe, two ports of the first bypass are respectively positioned at the upstream and the downstream of the second heat exchanger (4), and a first valve (13) is arranged on the first bypass; a second bypass is arranged on the sulfur-rich absorbent conveying pipeline, two ports of the second bypass are respectively positioned at the upstream and the downstream of the first heat exchanger (5), a second valve (12) is arranged on the second bypass, and the first valve (13) and the second valve (12) are respectively used for controlling the flow rate of the lean absorbent entering the second heat exchanger (4) and the flow rate of the sulfur-rich absorbent entering the first heat exchanger (5), so that the flash evaporation temperature of the sulfur-rich absorbent entering the flash tower (2) is controlled;
the flow ratio of the first bypass to the main road where the first bypass is located is 0-1: 3; the flow ratio of the second bypass to a main road where the second bypass is located is 0-1: 3;
the flash tower (2), the stripping tower (3) and the thermal regeneration tower (1) share the same tower body, the flash tower (2) is isolated from the stripping tower (3) through a first seal head (21), the stripping tower (3) is isolated from the thermal regeneration tower (1) through a second seal head (31), and the stripping tower (3) and the thermal regeneration tower (1) of the flash tower (2) are sequentially arranged from top to bottom; the liquid phase outlet of the flash tower (2) is connected with the inlet of the stripping tower (3) through an external pipeline; wherein the operating temperature in the flash tower (2) is 90-100 ℃, the flash pressure is 0.30-0.40 MPaG, the operating temperature in the gas stripping tower (3) is 90-100 ℃, the operating pressure is 0.30-0.40 MPaG, the operating temperature in the thermal regeneration tower (1) is 87-104 ℃, and the operating pressure is 0.22-0.28 MPaG.
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