US20240384192A1

US20240384192A1 - Process and apparatus for removing oxygenates from liquefied petroleum gas

Info

Publication number: US20240384192A1
Application number: US18/638,522
Authority: US
Inventors: Jeffrey Tyska; Benjamin Moore; David A. Wegerer; Susmitha Chakkingal; David Alejandro Roman Barreno; Javier Rios; William J. Whyman; Phuong T. M. Do
Original assignee: UOP LLC
Current assignee: Honeywell UOP LLC
Priority date: 2023-05-16
Filing date: 2024-04-17
Publication date: 2024-11-21
Also published as: CN121175397A; WO2024238341A1

Abstract

A process and apparatus for removing oxygenates from a petroleum stream of C3 and/or C4 hydrocarbons comprises water washing the petroleum stream to absorb oxygenates to provide a hydrocarbon stream lean in oxygenates and a water stream rich in oxygenates. The water stream is stripped to remove oxygenates into an oxygenate concentrated stream and an oxygenate-lean water stream. The lean water stream can be recycled to the water wash column.

Description

FIELD

The field is the purification of liquefied petroleum gas (LPG). The field may particularly relate to purifying light gas streams from a fluid catalytic cracking (FCC) unit.

BACKGROUND

Catalytic cracking can create a variety of products from larger hydrocarbons. Often, a feed of a heavier hydrocarbons, such as a vacuum gas oil, is charged to a catalytic cracking reactor, such as a fluid catalytic cracking (FCC) reactor. Various products may be produced, including a gasoline product and/or light product such as propylene and/or ethylene. As biorenewable feeds become more available, refiners are desirous of charging heavy biorenewable feeds to an FCC unit to crack them to motor fuels.
The oxygenate content of LPG from a FCC unit increases significantly when biorenewable, bio-oil, such as pyrolysis oil or vegetable oil, is co-processed with the fossil feed. LPG range oxygenates are known to cause issues in downstream processes such as catalytic polycondensation units, alkylation units and extractive mercaptan oxidation units, so the oxygenates must be removed to low levels to ensure smooth operation of the refinery and product that meets specification.
Oxygenate removal is one of the main challenges to coprocessing bio-oil in an FCC unit, which is often economically advantageous due to government subsidies or penalty avoidance.

BRIEF SUMMARY

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic drawing of a process and apparatus of the present disclosure.

FIG. 2 is a schematic drawing of an additional process and apparatus of the present disclosure.

DEFINITIONS

The term “communication” means that fluid flow is operatively permitted between enumerated components, which may be characterized as “fluid communication”.
The term “downstream communication” means that at least a portion of fluid flowing to the subject in downstream communication may operatively flow from the object with which it fluidly communicates.
The term “upstream communication” means that at least a portion of the fluid flowing from the subject in upstream communication may operatively flow to the object with which it fluidly communicates.
The term “direct communication” means that fluid flow from the upstream component enters the downstream component without passing through any other intervening vessel.
The term “indirect communication” means that fluid flow from the upstream component enters the downstream component after passing through an intervening vessel.
The term “bypass” means that the object is out of downstream communication with a bypassing subject at least to the extent of bypassing.
As used herein, the term “predominant” or “predominate” means greater than 50%, suitably greater than 75% and preferably greater than 90%.
The term “column” means a distillation column or columns for separating one or more components of different volatilities. Unless otherwise indicated, each column includes a condenser on an overhead of the column to condense and reflux a portion of an overhead stream back to the top of the column and a reboiler at a bottom of the column to vaporize and send a portion of a bottoms stream back to the bottom of the column. Feeds to the columns may be preheated. The top pressure is the pressure of the overhead vapor at the vapor outlet of the column. The bottom temperature is the liquid bottom outlet temperature. Overhead lines and bottoms lines refer to the net lines from the column downstream of any reflux or reboil to the column. Stripper columns may omit a reboiler at a bottom of the column and instead provide heating requirements and separation impetus from a fluidized inert media such as steam. Stripping columns typically feed a top tray and take main product from the bottom.
As used herein, the term “a component-rich stream” means that the rich stream coming out of a vessel has a greater concentration of the component than the feed to the vessel.
As used herein, the term “a component-lean stream” means that the lean stream coming out of a vessel has a smaller concentration of the component than the feed to the vessel.
As used herein, the term “separator” means a vessel which has an inlet and at least an overhead vapor outlet and a bottoms liquid outlet and may also have an aqueous stream outlet from a boot. A flash drum is a type of separator which may be in downstream communication with a separator that may be operated at higher pressure.
The term “Cx” is to be understood to refer to molecules having the number of carbon atoms represented by the subscript “x”. Similarly, the term “Cx−” refers to molecules that contain less than or equal to x and preferably x and less carbon atoms. The term “Cx+” refers to molecules with more than or equal to x and preferably x and more carbon atoms.

DETAILED DESCRIPTION

A once-through water wash may be used to remove oxygenates from an LPG stream. However, this process has significant drawbacks. The wash water rate needed to sufficiently decrease the oxygenate concentration is high, leading to high operating costs and high water makeup rates. The wash water becomes rich in oxygenates, requiring the capacity of the wastewater treatment system to increase substantially. Oxygenates also cause issues with performance of the wastewater treatment unit. The LPG range oxygenates are lost resulting in a lost opportunity for product value.
Use of adsorbent beds to further reduce the concentration of oxygenates may result in common oxygenates in the feed such as aldehydes poisoning the adsorbent. Once-through water washes are often insufficient to remove these oxygenates to levels required for economic usage of adsorbent. Removing oxygenates from the once-through water or adsorbent regeneration streams can be costly and energy intensive. Even if the oxygenates are extracted, their final destination can also pose issues. Some of the oxygenates may excessively build up if this stream is sent back to the conversion unit.
Hydrotreating the oxygenates or an oxygenate rich hydrocarbon stream can convert the oxygenates to hydrocarbons. However, hydrotreating these oxygenates will require significant amounts of hydrogen and generate a large exotherm in the hydrotreating reactor. Additionally, some of the oxygenates such as acetaldehyde and methanol would be hydrogenated to fuel gas range components as opposed to liquid products, minimizing the benefit of hydrogenation.
Other solvents such as methanol can be used for extraction, but they are often difficult to separate from the oxygenates themselves due to their boiling point similarity and the formation of azeotropes. Consequently, the extraction schemes can carry high capital and operational expense.
We propose the solution of bulk oxygenate removal in a water wash column followed by stripping oxygenates from the water stream. The washed LPG stream can be treated to remove acid gases and may undergo adsorption to further remove oxygenates to specified levels if necessary. A water stream stripped of oxygenates can be recycled back to the water wash column to drive down fresh-water usage. The concentrated oxygenate stripper stream may be oxidized in a thermal oxidizer to generate heat that can supply energy requirements. Because the oxygenates come from the bio-oil, the heat generated from the thermal oxidizer is from a renewable source.
The proposed solution efficiently removes oxygenates to avoid aggregation or complications in downstream units while eliminating wastewater generation. The heating value of the oxygenates generated from their thermal oxidation may be directly employed in the process or can be used to generate steam for usage elsewhere. The process and apparatus will enable coprocessing of at least about 5% pyrolysis oil or about 10% vegetable oil or higher such as about 15% pyrolysis oil or about 30% vegetable oil in an FCC unit, which has up until now not been economically feasible while maintaining low oxygenates in the LPG stream.
FIG. 1 depicts the process and apparatus 10 of the present disclosure. A petroleum stream in line 12 is provided. The petroleum stream may be a petroleum liquid stream such as an LPG stream comprising at least 90 wt % of C3 and/or C4 hydrocarbons. In a mixed C3/C4 stream either component may predominate or comprise about 20 to about 70 wt % of the mixed stream. The petroleum stream may comprise other hydrocarbons, such as about 0.5 to about 3 wt % C5 hydrocarbons.
The petroleum stream in line 12 may be an LPG stream from an FCC unit. Specifically, the LPG stream may comprise a liquid stream from an overhead receiver of a debutanizer column downstream of an overhead line from a FCC main column downstream of an FCC riser reactor.
The petroleum stream in line 12 is in the liquid phase. The petroleum steam may comprise about 250 to about 30,000 wppm of oxygenates. The petroleum stream laden with oxygenates is fed to a water wash column 14 to have the bulk of the oxygenates absorbed into a water stream. Two water streams may be fed to the water wash column 14. A fresh water, make-up, stream may be fed to the water wash column 14 at the highest feed location in line 16. A recycle water stream may be fed to the water wash column at an intermediate feed location in line 18. The petroleum stream in line 12 is fed to the water wash column 14 at the lowest feed location, so the water counter currently contacts the petroleum stream. The fresh water stream in line 16 and the recycle water stream in line 18 may have their feed locations reversed. The water wash column 14 may be in downstream communication with the petroleum stream in line 12, the fresh water stream in line 16 and the recycle water stream in line 18. The water wash column may include internal structures to facilitate contact between the water and the hydrocarbon streams.
In the water wash column 14 the water streams absorb oxygenates from the petroleum stream to produce a wash overhead stream in line 20 rich in hydrocarbons and an aqueous wash bottoms stream in line 22 rich in oxygenates. The wash overhead stream in line 20 is lean in oxygenates in an embodiment comprising about 25 to about 250 wppm oxygenates resulting in a bulk oxygenate removal from the hydrocarbon phase. The concentration of oxygenates in the wash overhead stream in line 20 will depend on the concentration of oxygenates in the petroleum stream in line 12. The wash water column 14 may also remove nitriles present in the petroleum stream in line 12. The wash water column 14 may separate the nitriles into the aqueous wash bottoms stream in line 22 which are further separated. Optionally, a bisulfide wash column 15 may be used to further remove oxygenates. Aqueous sodium bisulfide may be fed in line 24 to wash nitriles from the wash overhead stream in line 20. The bisulfide wash column 15 can reduce some specific oxygenates from the wash overhead stream in line 20 as well. The wash overhead stream may be taken in a bisulfide wash overhead stream in line 21 from the bisulfide wash column 15 lean in nitriles. An oxygenates rich stream may exit the bisulfide wash column in line 25.
The ratio of the recycle water stream to the petroleum stream in the wash water column 14 is typically about 1:1 to about 3:1 on a weight basis, and suitably about 3:2 to about 5:2 on a weight basis. The exact ratio depends on the oxygenate concentrations specified for downstream units and the speciation of oxygenates seen in the LPG stream. The ratio of the recycle water stream to the fresh water stream is typically about 25:1 to about 100:1 on a weight basis. The fresh water rate depends on the concentration of oxygenates left in the recycle water stream. The water wash column 14 may be operated at a temperature of about 20° C. to about 50° C. and a pressure of about 0.5 to about 2.3 MPa.
The aqueous wash bottom stream in line 22 extends from a bottom of the water wash column 14 and is heated by heat exchange with the recycle water stream in line 18. The wash bottom stream in line 22 is rich in oxygenates which must be removed to permit recycle of the water stream to the water wash column 14. Hence, the wash bottom stream is fed to an oxygenate stripper column 28 which is in downstream communication with the bottoms line 22 from the water wash column 14.
In the oxygenate stripper column 28, the wash bottoms stream in line 22 is stripped of volatile oxygenates to produce a stripper overhead stream in an overhead line 30 extending from an overhead of the stripper column concentrated in oxygenates and a stripped water stream lean in oxygenates in a bottoms line 32 extending from a bottom of the stripper column 28. Nitriles may also be separated in the stripper overhead stream in an overhead line 30. A reboil line 34 may be taken from the stripper bottoms line 32 to be reboiled in a stripper reboiler 33 and returned to the stripper column 28. The recycle stream in line 18 may also be taken from the stripper bottoms line 32 and pumped back to the water wash column 14 after it is cooled by heat exchange with the wash bottom stream in line 22 and further cooled in a recycle cooler.
The water wash column 14 may be in downstream communication with a bottoms line of the oxygenate stripper column 28 to utilize recycled water. The oxygenate content of the recycle water stream is typically no more than about 2.5 wt % and preferably no more than about 1 wt %. The stripper is run at a low overhead pressure, typically about 69 kPa(g) (10 psig) to about 138 kPa(g) (20 psig) and temperature of about 80° C. to about 120° C. which enables low pressure steam to be used as a heating medium in the reboiler. In an embodiment low pressure steam can be generated in the thermal oxidizer unit 40 and transported to the stripper reboiler 33 in a steam line 36 for providing the reboiler heating duty.
The stripper overhead stream in the stripper overhead line 30 concentrated in oxygenates may be transported to a thermal oxidizer unit 40 to combust oxygenates. Nitriles are also passed to the thermal oxidizer unit 40 with the stripper overhead stream in line 30. The stripper overhead line 30 may be devoid of a condenser to maintain the stripper overhead stream in the vapor phase to obviate revaporizing the stripper overhead stream in the thermal oxidizer unit 40. The water present in the stripper overhead line must be made up by the fresh water make up stream in line 16. The freshwater rate in the make-up stream in line 16 should match the rate of water lost in the overhead of the oxygenate stripper column 28 to ensure no oxygenate rich wastewater is generated.
The stripper overhead stream in line 30 may be heated by a heat exchange with a flue gas stream in a flue gas line 44 in a waste heat exchanger 42 and fed to a thermal oxidizer 46. The stripper overhead stream of concentrated oxygenates may be supplemented with oxygenate streams in lines 230 and 238 from a regenerator process and apparatus 200 in FIG. 2 . The thermal oxidizer 46 may be in downstream communication with the overhead line 30 of the oxygenate stripper column 28. A combustion air stream in line 48 which may be preheated and a fuel gas stream in line 50 may be supplied to the thermal oxidizer 46. The flow rate of the fuel gas stream in line 50 may be diminished or eliminated by the flow rate of oxygenates to the thermal oxidizer 46 in the stripper overhead stream. The inlet temperature to the thermal oxidizer 46 is typically in the range of 120° C. to about 500° C. with a pressure of about −1 kPa(g) to about 100 kPa(g). The outlet temperature is typically in the range of about 450 to about 1100° C. with a pressure of about −1 kPa(g) to about 50 kPa(g). The residence time in the thermal oxidizer 46 is between about 0.5 and about 2 seconds. Any suitable thermal oxidizer 46 could be used, including, but not limited to, an adiabatic thermal oxidizer chamber. The thermal oxidizer 46 can be forced draft, induced draft, or a combination of both. The thermal oxidizer 46 may include an inline burner.
In the thermal oxidizer 46, oxygenates are oxidized to water and carbon dioxide. The hydrogen sulfide and other sulfur compounds in the thermal oxidizer feed are oxidized to sulfur oxide, including SO2 and SO3, and water. A flue gas stream in the flue gas line 44 from the thermal oxidizer 46 comprises sulfur oxides (i.e., SO2 and SO3) and one or more of H2O, CO, CO2, NO₂, NO, N2 and O2. The flue gas stream in the flue gas line 44 is forwarded to the waste heat exchanger 42 for heat exchange with the concentrated oxygenate stream in line 30. The flue gas inlet temperature of the waste heat exchanger 42 is suitably in the range of about 450 to about 1100° C. with a pressure of about −2 kPa(g) to about 50 kPa(g). The flue gas outlet temperature is typically in the range of about 150 to about 425° C. with the same pressure range. The flue gas stream may also boil a steam stream in the steam line 36 by heat exchange in a heat recovery steam generator 43 to produce steam in line 36 perhaps a low pressure steam for the stripper reboiler 33. The steam generated in the steam exchanger 43 in line 36 may be used to reboil the stripper bottoms steam in the reboil line 34 and the regenerant reboiler 234 in FIG. 2 . Alternatively, a flue gas stream may be used to directly heat the process. In another aspect, a flue gas stream may be used to generate high pressure or medium pressure steam for export, followed by generation of low pressure steam for usage in the process.
A flue gas stream in line 44 from the waste heat exchanger 42 and/or the heat recovery steam generator 43 may also be used to preheat combustion air in a quench section 52. Air in line 54 may be fed to the air quench section 52 and indirectly or directly heat exchanged with the flue gas stream in line 44. Air in line 54 may be directly injected into the flue gas stream in line 44 as necessary to reduce the flue gas temperature. Alternatively, the combustion air stream in line 54 may be indirectly heat exchanged with the flue gas stream in line 44 to further cool the flue gas stream and heat the combustion air stream in line 54. The heated combustion air stream may then be transported in line 48 as combustion occurs in the thermal oxidizer 46. The inlet temperature of the flue gas stream to the quench section 52 is typically in the range of 150 to about 425° C. with a pressure of about −2 kPa(g) to about 50 kPa(g). The outlet temperature is typically in the range of about 150 to about 250° C. with a pressure of about −3 kPa(g) to about 50 kPa(g).
The quenched flue gas stream in line 56 must be treated to remove sulfur oxides. In an embodiment, a dry sorbent is injected from line 60 into the quenched flue gas stream in line 56. The dry sorbent may be pneumatically injected into the quenched flue gas stream. The sorbent injected into flue gas stream in line 56 may be charged to a dry sorbent reactor 62 to ensure sufficient mixing residence time is provided to achieve reaction of sufficient sulfur oxides to meet emissions requirements for the final vent gas stream or just to remove sufficient sulfur oxides to lower the acid dew point to a sufficient degree or to avoid sulfuric acid condensation altogether in the flue gas. Injecting the dry sorbent into the quenched flue gas stream in line 56 provides further heat recovery from the flue gas. By lowering the sulfuric acid condensation point, more heat can be extracted from the flue gas, resulting in a process with lower utility cost and lower greenhouse gas emissions.
The dry sorbent may comprise sodium or calcium adsorbents. Calcium adsorbent comprises calcium hydroxide and may react with sulfur oxides as in Formulas (1) and (2):
SO₂+Ca(OH)₂→CaSO₃·½H₂O+½H₂O (1),
and
SO₃+Ca(OH)₂→CaSO₄+H₂O (2).
Sodium adsorbent may comprise sodium carbonate (NaHCO₃) or Trona (Na₂CO₃·NaHCO₃·H₂O). The sodium sorbent is injected directly into the hot flue gas in which it is calcined into porous activated sodium carbonate (Na₂CO₃) as shown in Formula (3):
NaHCO₃→Na₂CO₃+CO₂+H₂O (3).
The thermal decomposition reaction of Formula (3) occurs rapidly at elevated temperatures such as 80 to about 800° C. The high surface area created enables fast gas-solid reactions between sulfur oxides and Na₂CO₃to form Na₂SO₄as shown in Formulas (4) and (5):
Na₂CO₃+SO₂+½O₂→Na₂SO₄+CO₂ (4), and
Na₂CO₃+SO₃+½O₂→Na₂SO₄+CO₂ (5).
The reaction of adsorbent and sulfur oxides produces a sulfate salt laden flue gas stream in line 64 at a temperature of about 260° C. to about 343° C. and a pressure of about −3 kPa(g) to about 50 kPa(g). The sulfate salt particles are solid and can be removed from the flue gas stream. Sulfate salt particles include CaSO ₃18 ½H₂O, CaSO₄or Na₂SO₄. The flue gas stream in line 64 may be cooled to below about 220° C. and fed to a particulate removal section 66. A sulfate salt residue stream may be removed from the particulate removal section 66 such as by an augur in line 68 and be sold as a valuable product for glass, detergent or paint manufacture.
In the particulate removal section 66 sulfate salts are separated from the sulfate depleted flue gas stream by either by fabric, steel or ceramic filters or electrostatic precipitators. For example, a bag-house filter may contain filter bags that allow for the clean flue gas to pass through while retaining the suspended solid particulates with periodic blow-back to clean particulates from the filters. Cooling of the sulfate salt flue gas stream in line 68 may not be necessary if high-temperature, ceramic or stainless steel filters are used in the particulate removal section 66. The flue gas stream in line 68 may be emitted to the stack.
If nitriles are in the concentrated oxygenate stream in the stripper overhead line 30, the thermal oxidizer 46 may convert the nitriles to NOx. A NOx reduction SCR unit which is not shown may be used to react ammonia with NOx generated in the thermal oxidizer 46 to produce molecular nitrogen and water. Any suitable NOx reduction catalyst can be used, including but not limited to, a ceramic, carrier material such as titanium oxide with active catalytic components such as oxides of base metals including TiO2, WO3 and V2O5, or an activated carbon-based catalyst.
The hydrocarbon-rich wash overhead stream in the bisulfide wash overhead line 21 from the bisulfide wash column 15 is oxygenate lean but still contains hydrogen sulfide and other sulfur compounds. These other sulfur compounds include carbonyl sulfide, disulfides and methyl sulfides that are not water extractable. Hence, the bisulfide wash overhead stream in line 21 may be introduced to a sulfur removal unit 70 to remove sulfur compounds. The sulfur removal unit 70 may comprise a mercaptan oxidation unit. Initially, the bisulfide wash overhead stream in line 21 may be introduced to an acid gas removal column 72 for absorbing hydrogen sulfide from the bisulfide wash stream in line 21. The acid gas removal column 72 may be in downstream communication with the overhead line 20 from the water wash column 14. Several different types of acid gas removal columns 72 can be used, including caustic washing, amine treatment, and sodium carbonate treatment units. In an exemplary embodiment, hydrocarbon-rich wash overhead stream in line 21 is washed with an alkaline stream in the acid gas removal column 72. The acid gas removal column 72 intimately mixes the hydrocarbon-rich wash overhead stream in line 21 with the alkaline stream in line 74, where the alkaline stream is an aqueous alkaline solution. If the alkaline solution is sodium hydroxide, the solution may be at a concentration of from about 5 wt % to about 20 wt % caustic in water. The caustic reacts with the hydrogen sulfide to produce sodium hydrosulfide and sodium sulfide, both of which are soluble in water and absorb into the alkaline stream in line 74. An amine may also be added to the alkaline solution if it is desired to remove carbonyl sulfide.
A desulfided wash stream in an overhead line 76 containing LPG hydrocarbons and a spent alkaline stream in a bottoms line 78 containing the alkaline material and sulfide reaction products, such as caustic and sodium hydrosulfide, exit the acid gas removal column 72. The alkali in the acid gas removal column 72 is gradually discharged and replaced with fresh alkali. Operating conditions for the acid gas removal column 72 are variable, but typically include ambient temperatures and pressures sufficient to keep the wash overhead stream 20 in the liquid phase. For example, temperatures from about 10° C. to about 60° C., and more typically about 30° C. to about 50° C. and pressures ranging from about 500 kPa to about 1.5 MPa can be used.
If hydrogen sulfide and sulfur concentration is desirably further reduced, the desulfided wash stream in line 76 may be treated in an extraction column 80 to remove any present or remaining mercaptans by reacting them with an alkaline stream such as caustic to produce mercaptan salts. The desulfided wash stream in line 76 is intimately contacted with an alkaline stream in line 82, where the desulfided wash stream in line 76 and the alkaline stream in line 82 are in the liquid phase. In an exemplary embodiment, the alkaline stream in line 82 is charged near a top of the extraction column 80, and the desulfided washed stream in line 76 is charged near a bottom of the extraction column 80. The concentration of the alkaline stream in line 82 varies, but typically ranges from about 10 to about 20 wt % caustic in water. The aqueous alkaline stream in line 82 does not form a solution or a suspension with the hydrocarbons in the desulfided washed stream in line 76, and the alkali is more dense than the hydrocarbons in the desulfided washed stream 76. Therefore, the alkali flows downwardly through the extraction column 80 as hydrocarbons in the desulfided washed stream in line 76 flow upwardly through the extraction column 80. In one embodiment, the extraction column 80 includes a plurality of trays configured to direct the heavier alkaline stream in line 82 through a tortuous path downwardly while the desulfided washed stream in line 76 is directed through a tortuous path upwardly, and the trays are designed to intimately mix and contact the two streams as they flow in a counter-current manner. In an alternate embodiment, the extraction column 80 includes packing or other structures to mix the alkali and hydrocarbons as they flow past each other. The extraction column 80 is sized to provide sufficient stages to react the mercaptans with the alkali, such as about 2 to about 6 stages or more. Exemplary operating conditions for the extraction column 80 include a temperature of about 10° C. to about 50° C. and a pressure sufficient to keep the washed feed stream in the liquid phase, such as about 500 kPa to about 1.5 MPa.
A mercaptan salt rich aqueous alkaline stream in a bottoms line 86 exits the extraction column 80 and includes the aqueous alkaline solution and mercaptan salts. An oxygen supply stream in line 90 is added to the mercaptan salt rich aqueous alkaline stream in line 86 to react with the mercaptan salts. In an exemplary embodiment, the oxygen supply stream 90 is air, but other oxygen-containing gases can also be used. Oxygen and water react with the mercaptan salts in a mercaptan oxidizer 94 to form disulfides and alkali. An unaided reaction rate is slow, and therefore an oxidation catalyst in line 96 is used to speed the oxidation reaction to produce the disulfides, and the oxidation catalyst in line 96 is added to the alkaline recirculation system on an as-needed basis. In an exemplary embodiment, the oxidation catalyst in line 96 is added to the rich alkaline stream in line 86 upstream of the mercaptan oxidizer 94, but the oxidation catalyst in line 96 could be added at other locations as well. Wash oil can also be added to the rich alkaline stream in line 86 to aid in the separation of alkaline and hydrocarbon to minimize disulfide content in the alkaline stream in line 82.
The oxidation catalyst in line 96 may be a metal chelate and can be in liquid or solid form. Several chelating agents can be used, such as phthalocyanines, tetraphenylporphyrins, or tetraphyidinoporphyrazines. Many chelating agents are not readily soluble in water, but water solubility can be increased by brominating, sulfonating, or carboxylating the chelating agents. The metal is one or more of iron, cobalt, manganese, molybdenum, or vanadium. In some embodiments, water soluble oxidation catalysts in line 96 are used, but insoluble forms of the oxidation catalyst in line 96 can be used in suspension or supported on a substrate that is either held in a fixed position in the mercaptan oxidizer 94 or maintained in a slurry with the alkaline stream. Suitable substrates include activated carbon, charcoal granules, thermoplastic polymers, exchange resins, and a wide variety of other materials. One exemplary oxidation catalyst in line 96 is iron phthalocyanine tetrasulfonate, but many other embodiments of an oxidation catalyst are possible.
The rich alkaline stream in line 86, including mercaptan salts, oxygen from the oxygen supply stream in line 90, and the oxidation catalyst in line 96 are heated and enter the mercaptan oxidizer 94. The mercaptan oxidizer 94 includes a packed bed 93, trays, or other structures that keep the aqueous alkaline solution and the water insoluble disulfides well mixed as the alkali flows through. The mercaptan salts are oxidized to disulfides, so essentially no mercaptans remain in a mixed alkaline/disulfide stream in line 98 exiting the overhead of the mercaptan oxidizer 94. Exemplary operating conditions for the mercaptan oxidizer 94 include a pressure of about 200 kPa to about 500 kPa (gauge) and a temperature of about 30° C. to about 60° C. The alkaline stream in line 82 is replenished with fresh alkali in line 97 as needed. The fresh alkaline stream in line 97 can be added in a wide variety of locations, including but not limited to the rich alkaline stream in line 86 upstream from the mercaptan oxidizer 94, as illustrated.
The mixed alkaline/disulfide stream in line 98 exits the mercaptan oxidizer 94 and enters the disulfide separator 100. The disulfide separator 100 has no agitation and has a sufficient volume to allow the water insoluble disulfides to separate from the aqueous alkaline solution. The mercaptan oxidizer 94 and the disulfide separator 100 work together to regenerate alkali and are fluidly coupled to the extraction column 80. In an exemplary embodiment, the disulfide separator 100 has a residence time of about of about 0.5 to about 3 hours. Any excess gases, such as excess nitrogen or oxygen from the oxygen supply stream in line 96, are vented from the disulfide separator 100 in a vent line 102. The vent line 102 can be directed to a scrubber or other pollution control device, and optionally includes a liquids entrainment separator (not illustrated) to prevent discharge of alkali or disulfides. The disulfide oil is less dense than the alkali, so the upper layer of disulfide oil exits near the top of the disulfide separator 100 in a disulfide stream in overhead line 104, and the alkaline stream in line 82 is recovered from near the bottom of the disulfide separator 100. The alkaline stream in line 82 contains small amounts of carryover disulfide, and these disulfides enter the extraction column 80 in the alkaline stream in line 82. The carryover disulfides are then combined with the hydrocarbons exiting the extraction column 80, because the disulfides are more soluble in the non-polar hydrocarbons than in the polar alkaline solution.
Mercaptans are removed from the desulfided washed stream in line 76 in the extraction column 80, and the hydrocarbons in the desulfided washed stream in line 76 exit the extraction column 80 in a treated wash stream in an overhead line 84. The hydrocarbons in the treated wash stream in line 84 also includes a low concentration of disulfides from the recovered alkaline stream in line 82. If the treated wash stream in the overhead line 84 comprising LPG hydrocarbons contains oxygenates over a specified level or if there are significant non-polar oxygenates like ethers, the treated wash stream may be forwarded to an oxygenate adsorption unit 110 to remove residual oxygenates that persist from the wash overhead stream in line 20. Residual disulfides in the hydrocarbon treated stream in line 84 are typical.
The oxygenate adsorption unit 110 contains an oxygenate adsorbent bed(s) 112 that is in downstream communication with an overhead line of said water wash column 14. The adsorbent in the adsorbent bed may be an alkali metal aluminosilicate which is able to remove oxygenates down to trace levels. The alkali metal may be sodium. A suitable adsorbent is ORG-E MOLSIV available from UOP LLC in Des Plaines, Illinois. The oxygenate adsorbent bed will remove oxygenates to a concentration below about 2 to about 30 wppm. Moreover, the adsorbent can adsorb oxygenates so that no single oxygenate may have a concentration more than about 1 to 8 wppm. An oxygenate depleted LPG stream may be provided in line 114 exiting the oxygenate adsorption unit 110. The oxygenate adsorption unit 110 may operate at temperatures from about 30° C. to about 50° C. and pressures ranging from about 500 kPa to about 1.5 MPa. The oxygenate depleted LPG stream in line 114 is a LPG product stream cleansed of oxygenates. It may be further processed to separate components to valuable products without concern for oxygenate impurities. The bisulfide wash column 15 removes some specific oxygenates from the wash overhead stream in line 20 and may also lead to a lesser quantity of absorbent in the downstream unit to achieve same oxygenate specifications in oxygenate depleted LPG stream in line 114.
FIG. 2 provides a process for regenerating the adsorbent bed 110 which does not show the connections of FIG. 1 . When the oxygenate adsorbent bed 112 in the oxygenate adsorption unit 110 is spent, it is taken off line for regeneration by disconnecting it from line 84 in the process and apparatus 10 of FIG. 1 . LPG is drained from the adsorbent bed 112 by valving to a LPG surge tank 115 through a down line 114 to ensure minimal loss of LPG and connected to the regeneration process and apparatus 200 of FIG. 2 by appropriate valving. Another adsorbent bed (not shown) may be connected to line 84 to continue with the process of oxygenate removal from LPG of FIG. 1 .
In FIG. 2 , a regenerant is pumped into the process and apparatus 200 in line 202. The regenerant may comprise an aliphatic hydrocarbon having 5 to 8 carbon atoms per molecule, preferably a normal hydrocarbon. Normal hexane is a suitable regenerant. Fresh make-up regenerant may be provided in line 203.
A regenerant in line 202 is pumped to a cooling line 204, a heating line 206 or a standby line 226. Three main “modes” of regeneration are included in the embodiment of FIG. 2 . For the heating mode the regenerant will flow through a plurality of heat exchangers as described later in detail before heading to the adsorption unit 110 while the cooling line 204, and the standby line 226 will be blocked with the valves thereon closed. In the cooling mode the regenerant will flow through a cooler 208 before heading to the adsorption unit 110 while the heating line 206 and the standby line 226 are blocked with the valves thereon closed. In the standby mode the regenerant will flow through standby line 226 back to a regenerant receiver 224 while the cooling line 204, and the heating line 206 are blocked with the valves thereon closed. All of the flow will be diverted into one of these three pathways as the regeneration cycle proceeds.
The first stage of regeneration is a hot regenerant stage. In the hot regenerant stage, an open valve on the heating line 206 directs the regenerant to feed heaters on the heating line 206 while the valve on cooling line 204 is closed. Three feed heaters may be provided including, in downstream order, a regenerant vaporizer 210, a regenerant superheater 212 and an electric superheater 214. The feed heaters 210, 212 and 214 vaporize the regenerant and superheat it to about 260° C. (500° F.) to about 316° C. (600° F.). A temperature sensor in the adsorbent bed 112 senses a temperature and compares it to a set point. If the temperature is lower than the set point, the electric superheater 214 provides additional duty. If the temperature is higher than the set point, the electric superheater 214 duty is reduced.
The regenerant from the heating line 206 is fed in a regenerant feed line 215 to the spent adsorbent bed 112 in the adsorption unit 110 to desorb oxygenates into the regenerant stream to fully regenerate the oxygenate adsorbent. During regeneration, the adsorbent unit 110 is operated at about 260° C. (500° F.) to about 316° C. (600° F.) and a pressure of about 200 kPa and about 500 kPa. The hot regenerant stream rich in desorbed oxygenates exits the adsorption unit 110 in the hot discharge line 216 through an open valve thereon while the valve on a cool discharge line 218 is closed. The hot regenerant rich in oxygenates is cooled in an adsorbent cooler 220 which may be an air cooler and perhaps a trim cooler 222 and enters the regenerant receiver 224.
After the adsorbent bed is regenerated in the hot regenerant stage, a cooling stage is initiated. The flow of the regenerant in line 202 is directed through the cooling line 204 and to the regenerant feed cooler 208 while the valve on the heating line 206 is closed. The bed is cooled by about 30° C. to about 50° C. in this manner. After cooling, a high-pressure nitrogen from line 250 is fed to the oxygenate adsorption unit 110 to drain regenerant from the oxygenate adsorbent bed 112. The drained regenerant is stored in a hexane surge drum 252, to avoid losses of the material. LPG is then slowly introduced into the adsorbent bed 112 perhaps in line 116 from the LPG surge drum 115 until the exotherm is minimized. The adsorbent bed 112 is then considered to be regenerated and may be put back online by connecting the oxygenate adsorption unit 110 to the hydrocarbon treated stream in line 84, perhaps in a lag configuration.
In the standby mode, regenerant in line 202 may also bypass the oxygenate adsorption unit 110 in line 226 and enter the regenerant receiver 224 after being cooled in a spillback cooler 221. Standby mode can be operative while the adsorbent bed 112 is between regeneration and adsorption operation stages.
The regenerant receiver 224 separates a hydrocarbon phase in a dried regenerant stream in an overhead line 228 from an aqueous phase including oxygenates desorbed from the adsorbent bed 112 in an oxygenated water stream in a bottoms line 230. The regenerant receiver 224 also serves to separate nitrogen from regenerant in the draining step. The oxygenated water stream in the bottoms line 230 may be transported to the thermal oxidizer unit 40 in FIG. 1 for combustion of the oxygenates. This ensures that no oxygenate rich wastewater stream is generated during normal operation, minimizing the effect to existing infrastructure. The dried regenerant stream in the overhead line 228 is fed to a regenerant column 232. The regenerant receiver 224 is operated at about 50° C. (122° F.) to about 80° C. (176° F.) and a pressure of about 25 kPa and about 150 kPa.
The regenerant column 232 may be in downstream communication with the oxygenate adsorbent bed 112 of the adsorption unit 110 during regeneration periods. In the regenerant column 232, oxygenates are stripped from the dried regenerant stream in the regenerant receiver overhead line 228. A hydrocarbon regenerated stream in a regenerant column overhead line 233 is condensed in a regenerant condenser 234 and separated in a regenerant column receiver 236 to produce a regenerant off gas stream in a net overhead line 238 and a reflux stream which is fed back to the column. The regenerant off gas stream in the net overhead line 238 comprises residual regenerant and oxygenates and may be fed to the thermal oxidizer 40 in FIG. 1 via line 30. A deoxygenated regenerant stream in a regenerant column bottoms line 240 may be split between a reboil stream in line 241 which is reboiled in a regenerant reboiler 243 and returned to the column and a recycle regenerant stream in line 242 which is recycled to regenerant line 202 and pumped back into the process. The reboiler 243 may be heated by steam from the thermal oxidizer unit 40.
In an embodiment, the petroleum stream may be depropanized to separate C3 hydrocarbons from C4 hydrocarbons, so that each stream could be water washed separately in a dedicated water wash column. This embodiment may operate to reduce water rates to the water wash columns if the bulk of the oxygenates are concentrated in either stream. In such an embodiment the water wash column bottoms streams could be processed in the disclosed process and apparatus together. However, the overhead C3 and C4 streams from each water wash column would be treated separately perhaps in duplicate processes to preserve the separation of C3 hydrocarbons from C4 hydrocarbons. In addition, any hydrogen sulfide in the feed may be concentrated in the C3 stream, allowing for the possibility for different metallurgy.
The foregoing process and apparatus provide an efficient way of removing oxygenates from a light hydrocarbon stream without generating vast volumes of waste water that must be treated. Indeed, no waste water requiring treatment is produced from the process and apparatus 10.

EXAMPLE

The process and apparatus of the present disclosure was simulated on LPG streams produced from co-processing the following feeds with VGO in an FCC unit. Oxygenate concentrations are provided in the LPG feed to wash column in Table 1, in the washed LPG stream from the wash column in Table 2, in the recycle water stream from the oxygenate stripper column Table 3 and from the adsorption unit in Table 4. Concentrations are in wppm.

TABLE 1

Concentration in LPG Feed to Water Wash Column

Compound/case	5% pyoil	15% pyoil	10% canola	30% canola

Dimethyl Ether	565	1695	0	0
Diethyl Ether	20	58	0	0
Acetaldehyde	100	300	340	1020
Propionaldehyde	650	1950	710	2130
Butyraldehyde	1	3	0	0
Methanol	6400	19200	0	0
Ethanol	5	15	7	20
Propanol	1	3	7	20
Acetone	380	1140	880	2640
Methyl Ethyl	50	150	0	0
Ketone
Total Oxygenates	8172	24514	1944	5830
H2S	1500	1500	N/A	N/A
Acetonitrile
	50	150	N/A	N/A
C3/C4 hydrocarbon	Balance	Balance	Balance	Balance

TABLE 2

Concentration in LPG from Water Wash Column

Compound/case	5% pyoil	15% pyoil	10% canola	30% canola

Dimethyl Ether

	104	65	0	0
Diethyl Ether	17	46	0	0
Acetaldehyde	0	0	1	0
Propionaldehyde	3	1	46	37
Butyraldehyde	0	0	0	0
Methanol	69	13	0	0
Ethanol	1	1	0	1
Propanol	0	1	0	0
Acetone	15	21	3	12
Methyl Ethyl	0	0	0	0
Ketone
Total Oxygenates	209	148	50	50
H2S	448	186	N/A	N/A
Acetonitrile
	20	8	N/A	N/A
Water	Balance	Balance	Balance	Balance
WW ratio,	4.1	6.5	2.3	3.3
Recycle:LPG

TABLE 3

Concentrations in Recycle Water to Wash Water Column

Compound/case	5% pyoil	15% pyoil	10% canola	30% canola

Dimethyl Ether

	112	241	0	0
Diethyl Ether	1	2	0	0
Acetaldehyde	24	44	144	303
Propionaldehyde	157	288	281	622
Butyraldehyde	0	0	0	0
Methanol	15625	24066	0	0
Ethanol	5	9	4	9
Propanol	0	0	3	7
Acetone	111	198	374	798
Methyl Ethyl	12	22	0	0
Ketone
Total Oxygenates	16047	24870	806	1739
H2S	143	125	N/A	N/A
Acetonitrile
	28	15	N/A	N/A
C3/C4 hydrocarbon	Balance	Balance	Balance	Balance

TABLE 4

Concentrations from Adsorption Unit

	Compound/case	5% pyoil	15% pyoil

Dimethyl Ether	<5	<5
Diethyl Ether	<5	<5
Acetaldehyde	<5	<5
Propionaldehyde	<5	<5
Butyraldehyde	<5	<5
Methanol	<5	<5
Ethanol	<5	<5
Propanol	<5	<5
Acetone	<5	<5
Methyl Ethyl Ketone	<5	<5
Total Oxygenates	<15	<15

The Total Oxygenates remaining are mainly ethers while the balance is other compounds.

SPECIFIC EMBODIMENTS

While the following is described in conjunction with specific embodiments, it will be understood that this description is intended to illustrate and not limit the scope of the preceding description and the appended claims.
A first embodiment of the disclosure is a process for removing oxygenates from a petroleum stream of C3 and/or C4 hydrocarbons comprising absorbing oxygenates from the petroleum stream into a water stream to produce a wash overhead stream rich in hydrocarbons and a wash bottoms stream rich in oxygenates; and stripping the wash bottoms stream to produce a stripper overhead stream concentrated in oxygenates and a stripped water stream lean in oxygenates. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph wherein the petroleum stream comprises fluid cracked product. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising recycling the stripped bottoms stream as the water stream. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising feeding make-up water to the absorbing step. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising thermally oxidizing oxygenates in the stripper overhead stream An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph recovering the energy via exchange with the process or steam generation An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising removing hydrogen sulfide from the wash overhead stream. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising removing nitriles from the wash overhead stream. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising adsorbing residual oxygenates from the wash overhead stream in a bisulfide-wash column. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising absorbing hydrogen sulfide with an alkaline solution to provide a desulfided wash stream and oxidizing mercaptans from the desulfided wash stream. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising charging the stripper overhead stream to the thermal oxidizing step without undergoing condensation. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising injecting dry sorbent to adsorb sulfur oxides from the flue gas from the thermal oxidizing step. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising regenerating an adsorbent loaded with residual oxygenates with an alkane feed comprising 5 to 8 carbons. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising heat exchanging the wash bottoms stream with the stripped bottoms stream.
A second embodiment of the disclosure is an apparatus for removing oxygenates from a petroleum stream comprising; a water wash column in communication with a petroleum stream and a water stream; an oxygenate stripper column in downstream communication with a bottoms line from the water wash column; and an acid gas removal column in downstream communication with an overhead line from the water wash column. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph further comprising a thermal oxidizer in downstream communication with an overhead line of the oxygenate stripper column. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph wherein the water wash column is in downstream communication with a bottoms line of the oxygenate stripper column to recycle water. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph further comprising an oxygenate adsorbent bed in downstream communication with an overhead line of the water wash column. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph further comprising a regenerant column in downstream communication with the oxygenate adsorbent bed during regeneration periods.
A third embodiment of the disclosure is a process for removing oxygenates from a petroleum stream of C3 and/or C4 hydrocarbons comprising absorbing oxygenates from the petroleum stream into a water stream to produce a wash overhead stream rich in hydrocarbons and a wash bottoms stream rich in oxygenates; stripping the wash bottoms stream to produce a stripper overhead stream concentrated in oxygenates and a stripped bottoms stream lean in oxygenates; and recycling the stripped bottoms stream to the absorbing step as the water stream. An embodiment of the disclosure is one, any or all of prior embodiments in this paragraph up through the third embodiment in this paragraph further comprising thermally oxidizing oxygenates in the stripper overhead stream.
Without further elaboration, it is believed that using the preceding description that one skilled in the art can utilize the present disclosure to its fullest extent and easily ascertain the essential characteristics of this disclosure, without departing from the spirit and scope thereof, to make various changes and modifications of the disclosure and to adapt it to various usages and conditions. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limiting the remainder of the disclosure in any way whatsoever, and that it is intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.
In the foregoing, all temperatures are set forth in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated.

Claims

1. A process for removing oxygenates from a petroleum stream of C3 and/or C4 hydrocarbons comprising:

absorbing oxygenates from the petroleum stream into a water stream to produce a wash overhead stream rich in hydrocarbons and a wash bottoms stream rich in oxygenates; and

stripping said wash bottoms stream to produce a stripper overhead stream concentrated in oxygenates and a stripped water stream lean in oxygenates.

2. The process of claim 1 wherein said petroleum stream comprises fluid cracked product.

3. The process of claim 1 further comprising recycling said stripped bottoms stream as said water stream.

4. The process of claim 1 further comprising feeding make-up water to said absorbing step.

5. The process of claim 1 further comprising thermally oxidizing oxygenates in said stripper overhead stream.

6. The process of claim 5 recovering the energy via exchange with the process or steam generation.

7. The process of claim 1 further comprising removing hydrogen sulfide from said wash overhead stream.

8. The process of claim 1 further comprising removing nitriles from said wash overhead stream.

9. The process of claim 6 further comprising adsorbing residual oxygenates from said wash overhead stream in a bisulfide-wash column.

10. The process of claim 6 further comprising absorbing hydrogen sulfide with an alkaline solution to provide a desulfided wash stream and oxidizing mercaptans from said desulfided wash stream.

11. The process of claim 5 further comprising charging said stripper overhead stream to the thermal oxidizing step without undergoing condensation.

12. The process of claim 5 further comprising injecting dry sorbent to adsorb sulfur oxides from the flue gas from the thermal oxidizing step.

13. The process of claim 7 further comprising regenerating an adsorbent loaded with residual oxygenates with an alkane feed comprising 5 to 8 carbons.

14. The process of claim 1 further comprising heat exchanging said wash bottoms stream with said stripped bottoms stream.

15. An apparatus for removing oxygenates from a petroleum stream comprising;

a water wash column in communication with a petroleum stream and a water stream;

an oxygenate stripper column in downstream communication with a bottoms line from said water wash column; and

an acid gas removal column in downstream communication with an overhead line from said water wash column.

16. The apparatus of claim 13 further comprising a thermal oxidizer in downstream communication with an overhead line of said oxygenate stripper column.

17. The apparatus of claim 13 wherein the water wash column is in downstream communication with a bottoms line of the oxygenate stripper column to recycle water.

18. The apparatus of claim 13 further comprising an oxygenate adsorbent bed in downstream communication with an overhead line of said water wash column.

19. The apparatus of claim 16 further comprising a regenerant column in downstream communication with said oxygenate adsorbent bed during regeneration periods.

20. A process for removing oxygenates from a petroleum stream of C3 and/or C4 hydrocarbons comprising:

absorbing oxygenates from the petroleum stream into a water stream to produce a wash overhead stream rich in hydrocarbons and a wash bottoms stream rich in oxygenates;

stripping said wash bottoms stream to produce a stripper overhead stream concentrated in oxygenates and a stripped bottoms stream lean in oxygenates; and

recycling said stripped bottoms stream to said absorbing step as said water stream.

21. The process of claim 19 further comprising thermally oxidizing oxygenates in said stripper overhead stream.