PROCESS FOR CONVERTING NAPHTHA TO PARAFFINS WITH HYDROCRACKED CHARGE STREAMS

Description

FIELD

The field is the conversion of naphtha to paraffins. The field may particularly relate to converting naphtha to ethane which is pyrolyzed.

BACKGROUND

Light olefin production is vital to the production of sufficient plastics to meet worldwide demand. Dehydrogenation is a process in which light paraffins such as ethane and propane can be dehydrogenated to make ethylene and propylene respectively, typically in the presence of a catalyst. Dehydrogenation can be achieved in either the presence of an oxidant such as oxygen or in the absence of an oxidant. Non-oxidative dehydrogenation is an endothermic reaction which requires external heat to drive the reaction to completion. Propane dehydrogenation (PDH) is a widely practiced example of non-oxidative dehydrogenation to produce propylene from propane. Ethane oxidative dehydrogenation is a newer oxidative process for converting ethane to ethylene which can be conducted at lower temperatures with lower carbon oxide emissions than non-oxidative and thermal cracking processes.

Fluid catalytic cracking (FCC) is another endothermic process that can be tuned to produce substantial propylene. However, not every FCC unit is tuned to make substantial propylene. Also, high propylene FCC units do not recover much ethylene; less than 1% of global ethylene supply comes from FCC.

The great bulk of the ethylene consumed in the production of plastics and petrochemicals such as polyethylene is produced by the thermal cracking of hydrocarbons. Steam is usually mixed with the feed stream to the cracking furnace to reduce the hydrocarbon partial pressure and enhance olefin yield and to reduce the formation and deposition of carbonaceous material in the cracking reactors. The process is therefore often referred to as steam cracking or pyrolysis.

Paraffins with a range of carbon numbers can be thermally cracked to produce olefins including ethane, propane, butanes, and naphtha. Ethane and naphtha feeds are typical due to higher light olefin yield than propane and butane feeds. Ethane feed is used in regions where light hydrocarbon gases are prevalent. In regions, where gas is not abundant, naphtha feed is employed for steam cracking. Naphtha steam cracking has long set the price in the ethylene industry due to higher production cost versus ethane steam cracking. The world does not currently produce enough ethane to supply the growing demand for ethylene. Therefore, regions lacking ethane supply such as Asia and Europe rely mainly on naphtha steam cracking to supply ethylene. Naphtha steam cracking yields only about 30%-35% ethylene with the balance including both relatively high-value by-products comprising propylene, butadiene, and butene-1 and relatively low value by-products comprising pyrolysis oil (pyoil), pyrolysis gasoline (pygas), and fuel gas. Additional pressures on naphtha steam cracking including minimum production requirements and environmental concerns have led to the withholding of government approvals in certain regions such as China. The ethylene industry needs a more efficient, economical and environmentally friendly route to light olefins from naphtha feeds.

BRIEF SUMMARY

A process for converting naphtha to paraffins is disclosed. The process comprises contacting a naphtha stream with a catalyst and hydrogen to produce a light paraffinic stream. The light paraffinic stream may be separated into an ethane stream and a propane stream. The ethane stream is thermally cracked to produce ethylene, propylene and a pyrolysis gasoline stream. The pyrolysis gasoline stream is hydrocracked to provide a cracked naphtha stream. The cracked naphtha stream may be converted to the ethane stream and the propane stream with the naphtha stream.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic drawing of a process for converting naphtha in accordance with an exemplary embodiment 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” or “predominance” means greater than 50%, suitably greater than 75% and preferably greater than 90%.

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.

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 “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.

DETAILED DESCRIPTION

In the proposed process, C3-C8+ hydrocarbon feedstock is first charged to a “Naphtha to Ethane and Propane” (NEP) unit to convert naphtha in the presence of hydrogen into desirable ethane and propane. The ethane produced is fed to an ethylene producing unit. The ethylene producing units provide over 75% yield of ethane to ethylene. The produced propane is fed to a propylene producing unit which provides over 85% yield of propane to propylene. The methane by-product from the NEP unit and the ethane and propane producing units can be used as a fuel including fuel needed to operate ethylene and propylene producing units which operate at elevated temperatures. Unconverted or under-converted C4+ components in the reactor outlet may be recycled for further processing to ethane and propane. The NEP process converts liquid feeds into ethane and propane, which may significantly increase the yields of ethylene and propylene. The NEP process also lowers the proportion of carbon dioxide produced per petrochemical product versus conventional processes. The disclosed NEP process provides for maximizing the production of ethylene and propylene from a complex comprising the NEP unit.

Turning to FIGURE, an embodiment of the process for converting naphtha 101 is disclosed. In an embodiment, the process 101 comprises a NEP reactor 120, an NEP separation unit 130, an ethane conversion unit 140, and a propane conversion unit 150. The process 101 also comprises a hydrocracking unit 110. A naphtha stream in line 111 may be combined with a hydrogen stream in line 131 to provide a charge stream in line 112. The charge stream in line 112 may be charged to the NEP reactor 120 to be contacted with an NEP catalyst. As described hereinafter in detail, the naphtha stream in line 111 may be taken from the hydrocracking unit 110. The naphtha stream may comprise C3 to C8+, perhaps C3 to C7, hydrocarbons preferably having a T10 between about −10° C. and about 60° C. and a T90 between about 70 and about 180° C. In an exemplary embodiment, the naphtha stream in line 111 may comprise C4 to C7 hydrocarbons. The naphtha feed stream may comprise normal paraffins, iso-paraffins, olefins, naphthenes, and aromatics.

In an aspect, the naphtha stream in line 111 may be heated before charging to the NEP reactor 111. In an embodiment, the naphtha stream may be heated to a reaction temperature of about 300° C. to about 600° C., suitably between about 325° C. and about 550° C., and preferably between about 350° C. and about 525° C. Overall, weight space velocity defined herein as hourly mass flow rate of feed in stream 111 divided by total catalyst in the NEP reactor 120 should be between about 0.3 to about 20 hr⁻¹, suitably between about 0.5 and about 10 hr⁻¹ and preferably between about 1 to about 4 hr⁻¹. A total pressure should be about 0.1 to about 3 MPa (abs), preferably greater than 1 MPa (abs). Under these conditions, C2-C3 yield is consistently in an excess of 80 wt %, while methane yield is less than about 16 wt %, suitably below about 14 wt % and typically below about 12 wt % and preferably no more than 10 wt %.

The hydrogen-to-hydrocarbon molar ratio is important to producing ethane and propane. The hydrogen-to-hydrocarbon ratio should be about 0.3 to about 15 and preferably about 0.5 to about 5. In a further embodiment, the hydrogen-to-hydrocarbon molar ratio may typically be no more than 5, suitably be no more than 3 and preferably be no more than 2. Low hydrogen-to-hydrocarbon ratio promotes desired reaction kinetics which are initiated with dehydrogenation. Hydrogen-to-hydrocarbon ratio may range from about 50% to about 500%, suitably no more than 300% and preferably no more than 200% of stoichiometric requirements to convert naphtha molecules to ethane and/or propane.

The NEP catalyst for converting naphtha to ethane and propane may contain a molecular sieve comprising large or medium pore mouths, that is, comprising 10 or 12 member rings, respectively. Examples of suitable molecular sieves include MFI, MEL, MFI/MEL intergrowth, MTW, TUN, UZM-39, IMF, UZM-44, UZM-54, MWW, UZM-37, UZM-8, UZM-8HS. Examples of suitable molecular sieves further include FER, AHT, AEL (SAPO-11), AFO (SAPO-41), MRE, MFS, EUO-1, TON (ZSM-22), MTT (ZSM-23) and UZM-53. Additional molecular sieves with larger pores include FAU, EMT, FAU/EMT intergrowth, UZM-14, MOR, BEA, UZM-50, MTW, ZSM-12. Additional examples include MSE and UZM-35.

MFI is a suitable NEP catalyst. It will be appreciated that ZSM-5 is an MFI-type aluminosilicate zeolite belonging to the pentasil family of zeolites and having a chemical formula of Na_nAl_nSi_96-nO₁₉₂·16H₂O (0<n<10). In various embodiments, the ZSM-5 zeolite may comprise a silica-to-alumina molar ratio of 20 to 1000, 20 to 800, 20 to 600, 20 to 400, 20 to 200 or 20 to 80. In various embodiments, the ZSM-5 zeolite may comprise a crystal size in the range of 10 to 600 nm, 20 to 500 nm, 30 to 450, 40 to 400 nm, or 50 to 300 nm.

The NEP catalyst may comprise a bound zeolite. The binder may comprise an oxide of aluminum, silicon, zinc, titanium, zirconium and mixtures of thereof. The binder may comprise a phosphate in the binder or a phosphate of the forenamed oxide binder materials. Preferably, the binder is a silicon oxide. The MFI zeolite may be supported in a silicon oxide containing binder or an alumina containing binder such as aluminum phosphate.

MFI zeolite slurry may be first mixed with a binder in the form of colloidal suspension (sol) and gelling reagent and then dropped into hot oil to make spheres controlled to produce ⅛-inch to about 1/32-inch diameter calcined supports. Alternatively, the zeolite may be mixed with a silicon oxide containing binder and extruded to 1/32 to ¼ inch diameter extrudates. Extrudates may be washed with ammonia to remove sodium ions from the zeolite, dried and calcined to remove the organic structural directing agent (OSDA) from the synthesized zeolite. Optionally, the calcined support may be ammonium-ion exchanged using an ammonium nitrate solution to remove residual sodium ions and dried at about 110° C.

The NEP catalyst comprises a metal on the catalyst. The metal may comprise a transition metal. In a further example, the metal may comprise platinum, palladium, iridium, rhenium, ruthenium and mixtures thereof. The metal may be a noble metal. A modifier metal may also be included on the catalyst. The modifier metal may include tin, germanium, gallium, indium, thallium, zinc, silver and mixtures thereof. The modifier metal should be more concentrated on the binder than on the zeolite. About 0.01 to about 5 wt % of each of the transition metal and the modifier metal may be on the catalyst.

Metal may be incorporated into the binder by evaporative impregnation. A solution of platinum such as tetraamine platinate nitrate or chloroplatinic acid may be contacted with the bound spherical or extrudate supports which have been calcined and ion-exchanged in a rotary evaporator, followed by drying and oxidation.

The NEP catalyst comprises a metal on the bound spherical or extrudate supports of the catalyst. Preferably, more of the metal is on the binder than on the zeolite. At least 60 wt %, suitably at least 70 wt %, preferably at least 80 wt % and most preferably at least 90 wt % of the metal is on the binder. The zeolite and/or the entire NEP catalyst is steamed oxidized to drive the metal off the zeolite. Steaming is preferably effected after the metal is added to the catalyst. The dried, impregnated spherical or extrudate supports may be steam oxidized in air for sufficient time to provide NEP catalysts. Steam oxidation in air at a temperature of about 500° C. to about 650° C. and about 5 mol % to about 30 mol % steam for about 1 to 3 hours may be suitable.

The NEP catalysts must be reduced to activate them for catalyzing the NEP reaction. For example, the catalyst may be reduced in flowing hydrogen at about 500 to about 550° C. for 3 hours before contacting feed.

After paraffin conversion, a light paraffinic stream is discharged from the NEP reactor 120 in an effluent line 122. The light paraffinic stream may comprise at least about 40 wt % ethane or at least about 40 wt % propane or at least about 70 wt % and preferably at least about 80 wt % ethane and propane. The ethane to propane ratio can range from about 0.1 to about 5. The light paraffinic stream can have less than about 16 wt %, suitably less than about 14 wt %, preferably less than about 12 wt %, and more preferably less than about 10 wt % methane.

The light paraffinic stream may be cooled and fed to an NEP separation unit 130. The NEP separation unit 130 may be a fractionation column or a series of fractionation columns and other separation units that may separate the light paraffinic stream in line 122 into the hydrogen stream in line 131, an ethane stream having a predominance of ethane in line 132, a propane stream having a predominance of propane in line 133 and a heavy aromatics stream in line 134. The NEP separation unit 130 may comprise a demethanizer column that separates the light paraffinic stream into a gas stream in an overhead line and a C2+ paraffin stream in a bottoms line. The gas stream may be sent to a hydrogen purification unit such as a PSA unit to recover hydrogen in line 131 for recycle to the NEP reactor 120 in a recycle line 131. Remaining methane from the hydrogen purification unit may be used for fuel gas. The C2+ paraffin stream may then be fed to a deethanizer column to produce the ethane stream in a deethanizer overhead line 132 and a C3+ paraffin stream in a deethanized bottoms line. The C3+ paraffin stream may then be fed to a depropanizer column to produce the propane stream in a depropanizer overhead line 133 and the heavy paraffin stream which may comprise C4+ hydrocarbons. The NEP separation unit 130 may take other forms.

For example, the NEP separation unit 130 may omit a demethanizer column and the light paraffinic stream in line 122 may feed a deethanizer column which produces a C2− stream in a deethanizer overhead line. The C2− stream can be separated in the hydrogen purification unit to recover a hydrogen stream in line 131 while residual ethane and methane from the hydrogen purification unit can comprise or supplement the ethane stream in line 132. The hydrogen purification unit may comprise a membrane unit and the hydrogen recovered from the membrane unit may be further purified in an absorption column before it is recycled to the NEP reactor 120 in lines 131. In an additional alternative, the C2− stream from the deethanizer column may be charged to an ethylene producing unit 140 in which ethane is converted to ethylene but methane and hydrogen rides through inertly to be recovered in a downstream ethylene recovery unit.

The ethane stream in line 132 may be charged to an ethylene producing unit 140 in which ethane in the ethane stream is converted into ethylene. In an embodiment, the ethylene producing unit 140 is a steam cracking unit. The ethane stream in line 132 may be cracked under steam in a furnace to produce a cracked stream including an ethylene stream 142. The ethane stream may be charged to the ethane steam cracking unit in the gas phase. The ethane steam cracking unit may preferably be operated at a temperature of about 750° C. (1382° F.) to about 950° C. (1742° F.). The cracked stream exiting the furnace of the ethane steam cracking unit may be in a superheated state. One or more quench columns, or other devices known in the art, but preferably an oil quench column and/or a water quench column, may be used for quenching or separating the cracked stream into a plurality of cracked streams. The ethane steam cracking unit may further comprise additional distillation columns, amine wash columns, compressors, expanders, etc. to separate the cracked stream into cracked streams rich in individual light olefins the most predominant of which is the ethylene stream in line 142. The ethylene stream may comprise a yield of at least 75 wt %, preferably at least 80 wt %, ethylene based on the ethane stream in line 132. Among the other components in the cracked stream exiting the ethane steam cracking, ethylene producing unit 140 may be hydrogen, methane, propylene, butene, and pyrolysis gasoline. Each of these components may be recovered and further processed.

The ethylene stream and a propylene stream from the ethylene producing unit 140 may be recovered or transported to polymerization plants, chemical plants or exported. Product recovery of at least 50 wt %, typically at least 60 wt % and suitably at least 70 wt % of valuable ethylene, propylene, and butylene products is achievable from the ethane steam cracking unit 140 based on the ethane stream in line 132.

Steam cracking may also generate less valuable by-products such as pyrolysis gasoline (pygas) and fuel oil (pyoil). The pyrolysis gasoline contains large proportions of paraffins and aromatics. The resulting paraffins include normal and non-normal paraffins which can be recovered or further processed. Aromatics are very stable and difficult to crack in a steam cracker. The present process comprises recycling this pyrolysis gasoline stream to recover paraffins to increase the yield of the process.

In accordance with the present disclosure, a pyrolysis gasoline stream is separated in line 143 from the ethylene producing unit 140. In an exemplary embodiment, the pyrolysis gasoline stream in line 143 may be hydrocracked to provide a naphtha stream.

The propane stream in line 133 may be charged to a propylene producing unit 150 in which propane in the propane stream is converted into propylene. The propylene producing unit 150 may be a propane dehydrogenation (PDH) unit. PDH catalyst is used in a dehydrogenation reaction process to catalyze the dehydrogenation of propane. The conditions in the dehydrogenation reactor may include a temperature of about 500 to about 800° C., a pressure of about 40 to about 310 kPa (abs) and a catalyst to oil ratio of about 5 to about 100.

The dehydrogenation reaction may be conducted in a fluidized manner such that gas, which may comprise the reactant paraffins with or without a fluidizing inert gas, is distributed to the reactor in a way that lifts the dehydrogenation catalyst in the reactor vessel while catalyzing the dehydrogenation of paraffins. During the catalytic dehydrogenation reaction, coke is deposited on the dehydrogenation catalyst leading to reduction of the activity of the catalyst. The dehydrogenation catalyst must then be regenerated in a regenerator. The regenerator may combust coke from the dehydrogenation catalyst and fuel gas to ensure sufficient enthalpy in the dehydrogenation reactor to promote the endothermic reaction.

The dehydrogenation catalyst selected should minimize cracking reactions and favor dehydrogenation reactions. Suitable catalysts for use herein include an active metal which may be dispersed in a porous inorganic carrier material such as silica, alumina, silica alumina, zirconia, or clay. An exemplary embodiment of a catalyst includes alumina or silica-alumina containing gallium, a noble metal, and an alkali or alkaline earth metal.

The catalyst support comprises a carrier material, a binder and an optional filler material to provide physical strength and integrity. The carrier material may include alumina or silica-alumina. Silica sol or alumina sol may be used as the binder. The alumina or silica-alumina generally contains alumina of gamma, theta and/or delta phases. The catalyst support particles may have a nominal diameter of about 400 to about 5000 micrometers with the average diameter of about 600 to about 3500 micrometers. Preferably, the surface area of the catalyst support is about 85 to about 140 m2/g.

The fluidized dehydrogenation catalyst may comprise a dehydrogenation metal on a support. The dehydrogenation metal may be a one or a combination of transition metals. A noble metal may be a preferred dehydrogenation metal such as platinum or palladium. Gallium is an effective metal for paraffin dehydrogenation. Metals may be deposited on the catalyst support by impregnation or other suitable methods or included in the carrier material or binder during catalyst preparation.

The acid function of the catalyst should be minimized to prevent cracking and favor dehydrogenation. Alkali metals and alkaline earth metals may also be included in the catalyst to attenuate the acidity of the catalyst. Rare earth metals may be included in the catalyst to control the activity of the catalyst. Concentrations of 0.001% to 10 wt % metals may be incorporated into the dehydrogenation catalyst. In the case of the noble metals, it is preferred to use about 10 parts per million (ppm) by weight to about 600 ppm by weight noble metal. More preferably it is preferred to use about 10 to about 100 ppm by weight noble metal. The preferred noble metal is platinum. Gallium should be present in the range of 0.3 wt % to about 3 wt %, preferably about 0.5 wt % to about 2 wt %. Alkali and alkaline earth metals may be present in the range of about 0.05 wt % to about 1 wt %.

Regenerated catalyst may be contacted with the propane stream in line 133 perhaps with a fluidizing gas to lift the propane stream and dehydrogenation catalyst up a riser while dehydrogenation occurs. Above the riser spent dehydrogenation catalyst and propylene product may be separated by a centripetal separation device. Propylene product gas may be quenched with a cooling fluid to prevent over reaction to undesired by-products. Separation of the propylene product from the PDH effluent stream in line 152 may include quench contacting and fractionation to produce a propylene product stream. Unreacted propane may be recycled to the dehydrogenation reactor and lighter gases may be recycled to the regenerator as fuel gas to be combusted to provide enthalpy for the reaction.

The propylene producing unit may also employ a catalytic moving bed reactor. The reactor section may comprise several radial flow reactors in parallel or series heated by charge and interstage heaters. The propane stream perhaps with added hydrogen flows in each dehydrogenation reactor from a screened center pipe through an annular dehydrogenation catalyst bed to an outer effluent annulus. Flow may be in the reverse fashion. The dehydrogenation catalyst may comprise a noble metal or mixtures thereof, a modifier selected from the group consisting of alkali metals or alkaline-earth metals and mixtures thereof, a component selected from the group consisting of tin, germanium, lead, indium, gallium, thallium, and mixtures thereof, and a porous support forming a catalyst particle. The catalyst support may comprise oil dropped alumina spheres.

Dehydrogenation conditions may include a temperature of from about 400 to about 900° C., a pressure of from about 0.01 to 10 atmospheres absolute, and a liquid hourly space velocity (LHSV) of from about 0.1 to 100 hr-1. The pressure in the dehydrogenation reactor is maintained as low as practicable, consistent with equipment limitations, to maximize chemical equilibrium advantages. Spent dehydrogenation catalyst in the annular catalyst bed may be withdrawn from the bottom of the bed, forwarded to a regenerator to combust coke from the catalyst with air at about 450 to about 600° C. Noble metal on the catalyst may be redispersed by an oxyhalogenation process, dried and returned to the top of the dehydrogenation catalyst bed as regenerated dehydrogenation catalyst.

Dehydrogenation effluent from the propylene producing unit 40 may be cooled, compressed, dried and hydrogen is cryogenically separated from the hydrocarbons with a net gas purity of 85 to 93 mol % hydrogen. Hydrocarbon liquid is selectively hydrogenated to convert diolefins and acetylenes and the hydrocarbon liquid is fractionated in a deethanizer column to remove ethane and propylene is split from propane in a propane-propylene splitter column to provide polymer-grade propylene. Propane may be recycled as feed to the propylene producing unit 150.

The heavy stream taken from the NEP separation unit 130 may be taken from a bottom of a depropanizer column in line 134. In an aspect, the heavy stream in line 134 may comprise greater than 98% aromatics. In accordance with the present disclosure, the heavy stream in line 134 may be hydrocracked to provide a naphtha stream.

In an embodiment, the pyrolysis gasoline stream in line 143 and the heavy stream in line 134 are fed to a hydrocracking reactor of the hydrocracking unit 110. A hydrocracking hydrogen stream in line 102 is also passed to the hydrocracking reactor. In an aspect, the hydrocracking hydrogen stream in line 102 may be taken from the hydrogen stream in line 131. The pyrolysis gasoline stream in line 143 and the heavy stream in line 134 are fed to a location in the lower half portion of the hydrocracking reactor. In an embodiment, the pyrolysis gasoline stream in line 143 is passed to a location above the heavy stream in line 134 in the lower half portion of the hydrocracking reactor.

In the hydrocracking unit 110, the pyrolysis gasoline stream in line 143 and the heavy stream in line 134 are cracked in the presence of hydrogen to lower molecular weight hydrocarbons comprising a hydrocracked product stream. Another feed stream in line 116 may also be fed to the hydrocracking unit 110. The hydrocracking unit 110 may comprise one or more fixed bed reactor(s) that comprises one or more vessels, single or multiple catalyst beds in each vessel, and various combinations of hydrotreating catalyst and/or hydrocracking catalyst in one or more vessels. If the hydrocracking unit 110 does not comprise a hydrotreating reactor, the catalyst beds in the hydrocracking reactor(s) may include a hydrotreating catalyst for the purpose of saturating, demetallizing, desulfurizing or denitrogenating the hydrocarbon feed stream before it is hydrocracked with the hydrocracking catalyst in subsequent vessels or catalyst beds in the hydrocracking unit. The hydrocracked product stream is taken in line 111 from the hydrocracking unit 110.

The hydrocracking catalyst may utilize amorphous silica-alumina bases or a crystalline zeolite cracking base upon which is deposited a Group VIII metal hydrogenating component. Additional hydrogenating components may be selected from Group VIB for incorporation with the amorphous silica alumina or zeolite base.

The zeolite cracking bases are sometimes referred to in the art as molecular sieves and are usually composed of silica, alumina and one or more exchangeable cations such as sodium, magnesium, calcium, rare earth metals, etc. They are further characterized by crystal pores of relatively uniform diameter between 4 and 14 Angstroms. It is preferred to employ zeolites having a relatively high silica/alumina mole ratio between 3 and 12. Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite and faujasite. Suitable synthetic zeolites include, for example, the B, X, Y and L crystal types, e.g., synthetic faujasite and mordenite. The preferred zeolites are those having crystal pore diameters between 8 and 12 Angstroms, wherein the silica/alumina mole ratio is 4 to 6. One example of a zeolite falling in the preferred group is synthetic Y molecular sieve.

The natural occurring zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic zeolites are nearly always prepared in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ion-exchanged with a polyvalent metal and/or with an ammonium salt followed by heating to decompose the ammonium ions associated with the zeolite, leaving in their place hydrogen ions and/or exchange sites which have actually been decationized by further removal of water. Hydrogen or “decationized” Y zeolites of this nature are more particularly described in U.S. Pat. No. 3,100,006.

By one approach, the hydrocracking conditions in the hydrocracking unit 110 may include a temperature from about 290° C. (550° F.) to about 468° C. (875° F.), preferably about 343° C. (650° F.) to about 445° C. (833° F.), a pressure from about 4.8 MPa (gauge) (700 psig) to about 20.7 MPa (gauge) (3000 psig), a liquid hourly space velocity (LHSV) from about 0.4 to less than about 2.5 hr⁻¹ and a hydrogen rate of about 421 Nm³/m³ (2,500 scf/bbl) to about 2,527 Nm³/m³ oil (15,000 scf/bbl).

In accordance with the present disclosure, the hydrocracked product stream is taken in line 111 is a cracked naphtha stream. The cracked naphtha stream in line is passed to the NEP reactor 120. In an exemplary embodiment, the hydrocracked product stream in line 111 comprises light naphtha having C4 to C7 hydrocarbons. The hydrocracked product stream in line 111 is combined with the hydrogen stream in line 131 to provide the charge stream in line 112. The charge stream in line 112 is fed to the NEP reactor 120 and processed as previously described.

The foregoing disclosure provides a process for converting naphtha with maximizing the yield of the process.

Example

Arab light crude was used as feedstock to the disclosed process. The process with a pyrolysis gasoline stream recycled to hydrocracking was compared with a process in which the pyrolysis gasoline stream was exported. The results are tabulated in the Table below:

As evident form the Table, 3758 kMTA of light olefins (2200 kMTA ethylene+1558 kMTA propylene) was produced from 4752 kMTA of Arab light feedstock at a yield of 79.1% for the pyrolysis gasoline export process. For the process with pyrolysis gasoline recycling back to the hydrocracking unit, the same 4752 kMTA of crude produced 3813 kMTA of light olefins (2200 kMTA ethylene+1613 kMTA propylene) for a yield of 80.2%. There was no export of the pyrolysis gasoline in the recycling process. Thus the production of light olefins (ethylene+propylene) was increased by recycling the pyrolysis gasoline (pygas) from the ethane steam cracking unit back to the hydrocracking unit.

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 converting naphtha comprising contacting a naphtha stream with a catalyst and hydrogen to produce a light paraffinic stream; separating the light paraffinic stream into an ethane stream and a propane stream; thermally cracking the ethane stream to produce a pyrolysis gasoline stream; and hydrocracking the pyrolysis gasoline stream to provide a cracked naphtha 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 producing an ethylene stream from cracking the ethane 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 separating the light paraffinic stream into an aromatics stream and hydrocracking the aromatics stream to provide the cracked naphtha 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 wherein the pyrolysis gasoline stream and the aromatics stream are hydrocracked in a hydrocracking reactor to produce the cracked naphtha 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 wherein the naphtha stream comprises light naphtha having C4 to C7 hydrocarbons. 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 step of thermally cracking the ethane stream comprises charging the ethane stream to a steam cracking unit and converting ethane in the ethane stream into ethylene. 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 hydrocracking step comprises contacting the pyrolysis gasoline stream with a hydrocracking catalyst comprising an amorphous silica-alumina or a crystalline zeolite and a metal hydrogenating component. 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 contacting the cracked naphtha stream and the naphtha stream with the catalyst and hydrogen to produce the light paraffinic 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 converting propane in the propane stream into propylene and charging the propane stream to a dehydrogenation unit. 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 hydrocracking a feedstock comprising the pyrolysis gasoline stream in a hydrocracking reactor to provide the cracked naphtha 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 separating the light paraffin stream into a hydrogen stream and recycling the hydrogen stream back to the contacting step.

A second embodiment of the disclosure is a process for converting naphtha comprising contacting a naphtha stream with a catalyst and hydrogen to produce a light paraffinic stream; separating the light paraffinic stream into an ethane stream and a propane stream; converting ethane in the ethane stream into ethylene and a pyrolysis gasoline stream; and hydrocracking a feedstock comprising the pyrolysis gasoline stream in a hydrocracking reactor to provide a cracked naphtha stream. 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 separating the light paraffinic stream into an aromatics stream and hydrocracking the aromatics stream in the hydrocracking reactor to provide the cracked naphtha stream. 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 pyrolysis gasoline stream and the aromatics stream are hydrocracked in a hydrocracking reactor to produce the cracked naphtha stream. 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 step of converting comprises charging the ethane stream to a steam cracking unit. 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 contacting the cracked naphtha stream and the naphtha stream with the catalyst and hydrogen to produce the light paraffinic stream. 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 naphtha stream comprises light naphtha having C4 to C8 hydrocarbons.

A third embodiment of the disclosure is a process for converting naphtha comprising contacting a naphtha stream with a catalyst and hydrogen to produce a light paraffinic stream; separating the light paraffinic stream into an ethane stream, a propane stream, and an aromatics stream; converting ethane in the ethane stream into an ethylene stream; and hydrocracking the aromatics stream to provide a cracked naphtha 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 cracking the ethane stream produce a pyrolysis gasoline stream and hydrocracking the pyrolysis gasoline stream to provide the cracked naphtha 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 contacting the aromatics stream and the pyrolysis gasoline stream with a hydrocracking catalyst in a hydrocracking reactor to produce the cracked naphtha 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 contacting the cracked naphtha stream and the naphtha stream with the catalyst and hydrogen to produce the light paraffinic 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.