Patent ID: 12203036

DETAILED DESCRIPTION OF THE DISCLOSURE

With regards to the hydrocarbon stream, it can contain a first diluent. In that case, said hydrocarbon stream contains at least 10 wt % of pyrolysis plastic oil. In a preferred embodiment, said hydrocarbon stream presents a bromine number of at most 150 g Br2/100 g, preferably at most 100 g Br2/100 g even more preferably at most 80 g Br2/100 g, the most preferred at most 50 g Br2/100 g as measured according to ASTM D1159. In a preferred embodiment, said hydrocarbon stream contains at least 25 wt % of pyrolysis plastic oil, preferably at least 50 wt % of pyrolysis plastic oil, even more preferably 75 wt % of pyrolysis plastic oil, in the most preferred embodiment at least 90 wt % of pyrolysis plastic oil. It is also possible to use pure pyrolysis plastic oil. In this latter case, the hydrocarbon stream is only pyrolysis plastic oil. The other component of said hydrocarbon stream may include any diluent able to limit the temperature increase at the first and/or second hydrotreating step. The diluent used for the first hydrotreating step (first diluent) can be the same or different as the diluent used for the second hydrotreating step. In other words, a diluent shall contain low amount, acceptable by a steam cracker or neither any olefins nor any diene. For instance, part of the purified hydrocarbon stream may be recycled and used as diluent. A naphtha can also be used as diluent. The use of naphtha as diluent is particularly advantaging. Indeed, in a preferred embodiment, the purified hydrocarbon stream is further sent to a steam cracker mixed together with a naphtha. The use of naphtha as diluent avoids further step of separation of the diluent. The effluent obtained at the end of the inventive process can then preferably be directly sent to the steam cracker. In a preferred embodiment, the pyrolysis plastic oil is diluted into naphtha having a boiling range from 15 to 250° C., preferably 38 to 150° C., as measured with method ASTM D2887 to form the hydrocarbon stream at a concentration of 50 wt %, preferably 75 wt % of pyrolysis plastic oil is diluted in naphtha, even more preferably 90 wt % of pyrolysis plastic oil is diluted in the naphtha.

With regards to the optional dewatering of the hydrocarbon stream, it consists in any method known in the art to remove the water present in a hydrocarbon stream. As non-limiting examples, water can be removed by decantation followed by separation. Water can also be removed in a flash drum. The hydrocarbon stream can alternatively or in addition to the other methods described, can be treated over a desiccant like an alumina or molecular sieve. The various method described above can be used independently or in any combination.

With regards to the optional desalting step, it consists in the desalting techniques known in the art. For instance, typical desalters comprise one or more tanks into which said hydrocarbon stream and water are added. The hydrocarbon stream and water are intensively mixed to enhance the phase interface, typically upstream of the settling tank. The salts from the hydrocarbon stream are extracted via the aqueous phase.

Desalting is a water-washing operation performed because of the negative effect of salts in the downstream processes due to scale formation, corrosion, and catalyst deactivation. These salts can be found in two forms: dissolved in emulsified water droplets in the pyrolysis plastic oil, as a water-in-oil emulsion, or suspended amorphous or crystalline solids. The negative effects of these salts in downstream processes are: salt deposit formation as scales where water is vaporized and corrosion by hydrochloric acid formation from hydrolysis of magnesium and calcium chlorides at high temperatures (about 350° C.) as follows:

Desalting involves mixing pyrolysis plastic oil and/or said hydrocarbon stream with washing water, using a mixing valve or static mixers to ensure a proper contact between the pyrolysis plastic oil and the water, and then passing it to a separating vessel, where separation between the aqueous and hydrocarbon phases is achieved. Since emulsions can be formed in this process, there is a risk of water carryover in the organic phase. In order to overcome this problem chemical demulsifiers are added to promote the emulsion breaking or an electric field across the settling vessel is applied to coalesce the polar salty water droplets, and, therefore, separation of salty water is achieved.

In order to enhance the effective mixing between the hydrocarbon and aqueous phases and ensure the proper extraction of the salts and minerals into the aqueous phase, a mixing valve is used over which a pressure drop result in shear stress over the droplets that promotes an intimate water and oil contact. In addition to the mixing valve, upstream premixing devices can be used, such as spray nozzles or static mixers. The shear stress needs to be optimized to reach the right balance between smaller droplets, which improves the contact among the phases but however result in more stable emulsion.

Subsequently, the mixture goes to the desalter, a horizontal cylindrical tank that provides long enough residence time to separate the water and oil mixture in two phases. Some water droplets diameters are so small that they do not separate by gravity; so, an electrostatic field between two electrodes installed into the desalter is used to promote coalescence.

When emulsion is too stable and break only slowly, demulsifiers can be used. Demulsifiers are surfactants, when present at low concentration, interacts with the interfaces of the system, altering the interfacial free energies of those interfaces. In particular lipophilic anionic surfactants can be expected in the pyrolysis plastic oil, resulting from the presence of polyesters, polycarbonate and polyamides and from the presence of additives like, antioxidants and UV stabilizers, containing phenols, other oxygenated aromatics and phosphorus containing compounds and slip agents, like fatty acid amides, fatty acid esters, metallic stearates (for example, zinc stearate). An emulsion breaker will lower the interfacial tension and result in coalescence. A hydrophilic demulsifiers will balance the lipophilic surfactant.

In an embodiment, the demulsifying agent can be chosen among water, steam, acids, caustic solutions, complexing agents and their mixtures. Acids are for example strong acids, in particular inorganic acids, such as phosphoric acid, sulphuric acid. Complexing agents are for example weak organic acids (or their corresponding anhydrides) such as acetic acid, citric acid, oxalic acid, tartaric acid, malic acid, maleic acid, fumaric acid, aspartic amino acid, ethylenediaminetetraacetic acid (EDTA). Preferably, the demulsifying agent comprises water, steam, phosphoric acid, acetic acid, citric acid, oxalic acid, tartaric acid, malic acid, fumaric acid, aspartic amino acid, ethylenediaminetetraacetic acid, alkali, salts, chelating agents, crown ethers, or maleic anhydride.

With regards to traps for the silicon and/or metals and/or phosphorous and/or halogenates, it consists in silica gel, clays, alkaline or alkaline earth metal oxide, iron oxide, ion exchange resins, active carbon, active aluminium oxide, molecular sieves, and/or porous supports containing lamellar double hydroxide modified or not and silica gel, or any mixture thereof used in the fixed bed techniques known in the art. The trap is able to capture silicon and/or metals and/or phosphorous and/or halogenates, being preferably chosen among Ca Mg Hg via absorption and/or adsorption or it can also be constituted of one or more active guard bed with an adapted porosity. It can work with or without hydrogen coverage. The trap can be constituted of an adsorbent mass such as for instance a hydrated alumina. Molecular sieves can also be used to trap silicon. Other adsorbent can also be used such as silica gel for instance. The silicon trap is preferably able to trap organic silicon. Indeed, it is possible that the silicon present in the streams are in the form of organic silicon.

In a preferred embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with activated carbon. Activated carbon possesses preferably a high surface area (600-1600 m2/g), and is preferably porous and hydrophobic in nature. Those properties lead to a superior adsorption of non-polar molecules or little ionized molecules. Therefore, activated carbon can be used to reduce for instance siloxane from the liquid feed at temperature from 20 to 150° C., at pressures from 1 to 100 bar or from vaporized feed from 150 to 400° C. at pressure from 1 to 100 bar. Regeneration of saturated adsorbent can be performed via heating while using a sweeping gas.

In a preferred embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with silica or silica gel. Silica gel is an amorphous porous material, the molecular formula usually as (SiO2).nH2O, and unlike activated carbon, silica gel possesses polarity, which is more conductive to the adsorption of polar molecules. Because of —Si—O—Si— bonds, siloxanes exhibit partial polar character, which can contribute to adsorb on silica gel surface. The adsorption force of silica gel is often weak enough allowing regeneration of silica gel by heat treatment above 150 up to 300° C. using a sweeping gas.

In a preferred embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with molecular sieves. Molecular sieves are hydrous aluminosilicate substance, with the chemical formula Na2O·Al2O3.nSiO2.xH2O, which possesses a structure of three-dimensional crystalline regular porous and ionic exchange ability. Compared with silica gel, molecular sieves favour adsorption of high polarity. The regeneration of exhausted absorbents can be achieved via heating at high temperature to remove siloxane. Often, the regeneration is less efficient as the siloxanes might react irreversibly with the molecular sieve. In a most preferred embodiment, the molecular sieves are ion-exchanged or impregnated with a basic element such as Na. Na2O impregnation levels range from 3-10% wt typically and the type of sieve are typically of the A or faujasite crystal structure.

In a preferred embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with activated aluminium oxide. Activated aluminium oxide possesses large surface area (100-600 m2/g), which shows high affinity for siloxanes but also for polar oxide, organic acids, alkaline salts, and water. It can be an alkaline or alkaline-earth or rare-earth containing promoted alumina, the total weight content of these doping elements being less than 20% wt, the doping elements being preferably selected from Na, K, Ca, Mg, La, or mixture thereof. It can also be a metal promoted alumina where the metal is selected from group VI-B metal with hydrogenating activity such as Mo, W and/or from group VIII metal, such as Ni, Fe, Co

In another embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with alkaline oxide. Alkaline oxide for high temperature treatment such as calcium oxide (CaO) has strong activity to breakdown siloxanes and can be used as non-regeneratable adsorbent at temperature between 150 and 400° C.

In another embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with porous supports containing lamellar double hydroxides, being preferably an hydrotalcite. The hydrotalcite can comprise one or more metals with hydrogenating capacity selected from group VIB or Group VIII, preferably Mo. Those metals can be supported on the surface of the hydrotalcite, or can have been added to the actual structure of the lamellar double hydroxide, in complete or partial substitution, as an example, but without limiting the scope of the present invention, the divalent metal, usually Mg, can be exchanged for Ni, or the trivalent metal, substituted by Fe instead of Al.

The above-mentioned solid adsorbents can be used alone or in any combination in order to optimize the removal of silicon and/or metals and/or phosphorous and/or halogenates.

In another embodiment, silicon and/or metals and/or phosphorous and/or halogenates are trapped with a multi layered guard bed comprising at least two layers wherein the layer on the top of the bed is selected from silica gel, clays, alkaline or alkaline earth metal oxide, iron oxide, ion exchange resins, active carbon, active aluminium oxide, molecular sieves and wherein the layer on the bottom of the bed is selected from silica gel, clays, alkaline or alkaline earth metal oxide, iron oxide, ion exchange resins, active carbon, active aluminium oxide, molecular sieves. More preferably said layer on the top of the guard bed comprises silica gel and/or active carbon and said layer on the bottom of the guard bed comprises molecular sieves and/or active aluminium oxide.

In another embodiment, when the pyrolysis plastic oil contains high quantities of HCl and/or Halogenated compounds (namely at least 500 ppm wt of HCl based on the total amount of pyrolysis plastic oil), particular adsorbents can be used such as silica, clays—such as bentonite, hydrotalcite—alkaline or alkaline earth metal oxide—such as iron oxides, copper oxides, zinc oxide, sodium oxide, calcium oxide, magnesium oxide-alumina and alkaline or alkaline-earth promoted alumina-, iron oxide (hematite, magnetite, goethite), ion exchange resins or combination thereof. In a most preferred embodiment, silicon and/or metals and/or phosphorous and/or halogenates containing at least 500 ppm wt of HCl based on the total amount of pyrolysis plastic oil are trapped with activated alumina. As HCl is a polar molecule, it interacts with polar sites on the alumina surface such as hydroxyl groups. The removal mechanism relies predominantly on physical adsorption and low temperature and the high alumina surface area is required to maximize the capacity for HCl removal. The HCl molecules remain physically adsorbed as a surface layer on the alumina and can be removed reversibly by hot purging. Promoted aluminas are a hybrid in which a high alumina surface area has been impregnated with a basic metal oxide or similar salts, often of sodium or calcium. The alumina surface removes HCl through the mechanisms previously described, however the promoter chemically reacts with the HCl giving an additional chloride removal mechanism referred to as chemical absorption. Using sodium oxide as an example of the promoter, the HCl is captured by formation of sodium chloride. This chemical reaction is irreversible unlike physical adsorption and its rate is favoured by higher temperature. The promoted alumina chloride guards are very effective for liquid feeds due to the irreversible nature and high rate of the chemical reaction once the HCl reaches the reactive site.

Another class of chemical absorbents combines Na, Zn and Al oxides in which the first two react with HCl to form complex chloride phases, for example Na2ZnCl4and the chemical reactions are irreversible. U.S. Pat. Nos. 4,639,259 and 4,762,537 relate to the use of alumina-based sorbents for removing HCl from gas streams. U.S. Pat. Nos. 5,505,926 and 5,316,998 disclose a promoted alumina sorbent for removing HCl from liquid streams by incorporating an alkali metal oxide such as sodium in excess of 5% by weight on to an activated alumina base. Other Zn-based products range from the mixed metal oxide type composed of ZnO and Na2O and/or CaO. The rate of reaction is improved with an increase in reactor temperature for those basic (mixed) oxides.

With regards to the optional guard bed to trap solid particles located on the top of said first and/or second hydrotreating, it is placed on the top of said first and/or second hydrotreating to remove the solid particles remaining in the feed such as coke particles coming from heating tubes, iron scales from corrosion, dissolved impurities such as iron, arsenic, calcium-containing compounds, sodium chloride, silicon contained in upstream additives, etc. Grading materials which have high void space to accumulate and ‘store’ these particulates are frequently used. Effective feed filtration to remove particulates in combination with high void grading provides a longer mitigation of pressure drop buildup. In a preferred embodiment, said guard bed to trap solid particles has a continuously decreasing particle size including a region 25 to 150 centimeters of particles, having a fraction of 0.3 to 2.0 cm diameter range. Since such guard beds to trap solid particles are designed specifically to handle the contaminants, they help to prolong the life of the hydrotreating catalyst and require fewer total catalyst changeouts.

With regards to the impurities removal treatment step to remove silicon, phosphorous, metals and/or halogenate compounds, it consists preferably of a solvent extraction unit. The solvent can be water, alcohol, NaOH, KOH, etc. For example, the silicon extraction with NaOH described in the COMET patent (EP2643432B1), the metals solvent extraction unit used in the refining of used oils.

With regards to the first hydrotreating step, it consists mainly in the hydrogenation phase to saturate the conjugated diene and alkynes into mainly olefins. Depending on the composition of the hydrocarbon stream, the first step hydrotreating is performed either in liquid phase or in trickle bed mode. This step is well known in steam cracking unit as 1ststep hydrogenation of pyrolysis gasoline. The first hydrotreating step will hydrogenate the diene and in particularly the conjugated diene and acetylenic bonds. The first hydrotreating step will lead to a decrease the diene value. The decrease of the diene value observed between the inlet and the outlet of the first hydrotreating step should be of at least 10% preferably at least 25% as measured according to UOP 326.

With regards to the second hydrotreating step, it consists in a step at a temperature higher than 200° C., in presence of hydrogen with well-known catalysts to hydrogenate the olefins and to convert sulfur, nitrogen components into respectively H2S and NH3. Depending on the composition of the stream entering this second hydrotreating step, it is either performed in gas phase or the reactor operates in trickle bed mode. This step can have also a metal trap function, a cracking function, a de-aromatization function depending of the characteristic of the catalyst and the used operating condition. This step can be performed in one reactor with different layers of catalysts or several reactors in series depending of the function sought.

In a preferred embodiment, said second hydrotreating step is performed over at least one catalyst that presents both (i) an hydrotreating function, and (ii) a trap function. In that case, the preferred operating conditions advantageously be the following: the preferred inlet temperature is of at least 200° C. and at most 500° C.; the preferred LHSV is between 1 to 10 h−1, preferably 2 to 4 h−1; the preferred pressure ranges from 10 to 90 barg in presence of H2; the ratio H2/hydrocarbon ranges from 200 NL/L to 900 NL/L, preferably in the presence of at least 0.005 wt %, preferably 0.05 wt % even more preferably 0.5 wt % of sulphur, being preferably H2S or organic sulfur compounds, in the stream. The use of such catalyst is particularly advantaging because it allows to simultaneously perform the hydrotreating reaction and to trap impurities like silicon that may still be present in the stream.

With regards to the waste plastic pyrolysis, an example of a pyrolysis process for waste plastics is disclosed in U.S. Pat. No. 8,895,790 or in US2014/0228606 and in WO 2016/009333.

In a waste plastic pyrolyzer, mixed plastics (e.g., waste plastics) are placed in pyrolysis unit or pyrolyzer. In the pyrolysis unit, the waste plastic is converted via pyrolysis to a pyrolysis product, wherein the pyrolysis product comprises a gas phase (e.g., pyrolysis gases, such as C1 to C4 gases, hydrogen (H2), carbon monoxide (CO), carbon dioxide (CO2) mainly) and a liquid phase being pyrolysis plastic oil. The plastic waste may include post-consumer waste plastics, such as mixed plastic waste. Mixed plastics can comprise non-chlorinated plastics (e.g., polyolefins, polyethylene, polypropylene, polystyrene, copolymers, etc.), chlorinated plastics (e.g., polyvinylchloride (PVC), polyvinylidene chloride (PVDC), etc.), and the like, or mixtures thereof. Generally, waste plastics comprise long chain molecules or polymer hydrocarbons. Waste plastics may also include used tires.

The pyrolysis unit may be any suitable vessel configured to convert waste plastics into gas phase and liquid phase products (e.g., simultaneously). The vessel may be configured for gas phase, liquid phase, vapor-liquid phase, gas-solid phase, liquid-solid phase, or slurry phase operation. The vessel may contain one or more beds of inert material or pyrolysis catalyst comprising sand, zeolite, alumina, a catalytic cracking catalyst, or combinations thereof. Generally, the pyrolysis catalyst is capable of transferring heat to the components subjected to the pyrolysis process in the pyrolysis unit. Alternatively, the pyrolysis unit can be operated without any catalyst (e.g., pure thermal pyrolysis). The pyrolysis unit may be operated adiabatically, isothermally, nonadiabatically, non-isothermally, or combinations thereof. The pyrolysis reactions of this disclosure may be carried out in a single stage or in multiple stages. For example, the pyrolysis unit can be two reactor vessels fluidly connected in series.

In a configuration where the pyrolysis unit comprises two vessels, the pyrolysis process may be divided into a first stage which is performed in a first vessel and in a second stage fluidly connected downstream of the first stage which is performed in the second vessel. As will be appreciated by one of skill in the art, and with the help of this disclosure, the second stage may enhance the pyrolysis of an intermediate pyrolysis product stream flowing from the first stage into the second stage, to yield a pyrolysis product flowing from the second stage. In some configurations, the first stage may utilize thermal cracking of the waste plastics, and the second stage may utilize thermal or catalytic cracking of the waste plastics to yield the pyrolysis product flowing from the second stage. Alternatively, the first stage may utilize catalytic cracking of the waste plastics, and the second stage may utilize thermal or catalytic cracking of the waste plastics to yield the pyrolysis product flowing from the second stage.

In some configurations, the pyrolysis unit may include one or more equipment configured to convert mixed plastics into gas phase and liquid phase products. The one or more equipment may or may not contain an inert material or pyrolysis catalyst as described above. Examples of such equipment include one or more of heated extruders, heated rotating kiln, heated tank-type reactors, packed bed reactors, bubbling fluidized bed reactors, circulating fluidized bed reactors, empty heated vessels, enclosed heated surfaces where plastic flows down along the wall and cracks, vessels surrounded by ovens or furnaces, or any other suitable equipment offering a heated surface to assist in cracking.

The pyrolysis unit can be configured to pyrolyze (e.g., crack), and in some aspect (e.g., where hydrogen is added to the pyrolysis unit), additionally hydrogenate components of the waste plastic stream fed to the pyrolysis unit. Examples of reactions which may occur in the pyrolysis unit include, but are not limited to isomerization of one or more normal paraffins to one or more i-paraffins, selective ring opening of one or more cycloparaffins to one or more i-paraffins, cracking of long chain length molecules to short chain length molecules, removal of heteroatoms from heteroatom-containing hydrocarbons (e.g., dechlorination), hydrogenation of coke generated in the process, or combinations thereof.

In one or more configurations of the pyrolysis unit, a head space purge gas can be utilized in all or a portion of the pyrolysis stage(s) (conversion of waste plastics to a liquid phase and/or gas phase products) to enhance cracking of plastics, produce valuable products, provide a feed for steam cracking, or combinations thereof. The head space purge gas may include hydrogen (H2), C1 to C4 hydrocarbon gases (e.g., alkanes, methane, ethane, propane, butane, isobutane), inert gases (e.g., nitrogen (N2), argon, helium, steam), and the like, or combinations thereof. The use of a head space purge gas assists in the dechlorination in the pyrolysis unit, when the waste plastic comprises chlorinated plastics. The head space purge gas may be introduced to the pyrolysis unit to aid in the removal of volatiles entrained in the melted mixed plastics present in the pyrolysis unit.

A hydrogen (H2) containing stream can be added to the pyrolysis unit to enrich the pyrolysis unit environment with H2, assist in stripping entrapped hydrogen chloride in the pyrolysis unit, provide a local environment rich in hydrogen in the pyrolysis melt or liquid, or combinations thereof; for example via a H2 containing stream fed directly to the pyrolysis unit independently of the waste plastic stream. In some aspects, H2 can also be introduced along with stream to the pyrolysis unit, with adequate safety measures incorporated for hydrogen handling with plastics feed.

The pyrolysis unit may facilitate any reaction of the components of the waste plastic stream in the presence of, or with, hydrogen. Reactions may occur such as the addition of hydrogen atoms to double bonds of unsaturated molecules (e.g., olefins), resulting in saturated molecules (e.g., paraffins, i-paraffins, naphthenes). Additionally or alternatively, reactions in the pyrolysis unit may cause a rupture of a bond of an organic compound, with a subsequent reaction and/or replacement of a heteroatom with hydrogen.

The use of hydrogen in the pyrolysis unit can have beneficial effects of i) reducing the coke as a result of cracking, ii) keeping the catalyst used (if any) in the process in an active condition, iii) improving removal of chloride from stream such that the pyrolysis product from pyrolysis unit is substantially dechlorinated with respect to waste plastic stream, which minimizes the chloride removal requirement in units downstream of the pyrolysis unit, iv) hydrogenating of olefins, v) reducing diolefins in pyrolysis product, vi) helping operate the pyrolysis unit at reduced temperatures for same levels of conversion of waste plastic stream in the pyrolysis unit, or combinations of i)-vi).

The pyrolysis processes in the pyrolysis unit may be low severity or high severity. Low severity pyrolysis processes may occur at a temperature of less than about 450° C., alternatively 250° C. to 450° C., alternatively 275° C. to 425° C., or alternatively 300° C. to 400° C., and may produce pyrolysis oils rich in mono- and di-olefins as well as a significant amount of aromatics. High severity pyrolysis processes may occur at a temperature of equal to or greater than about 450° C., alternatively 450° C. to 750° C., alternatively 500° C. to 700° C., or alternatively 550° C. to 650° C., and may produce pyrolysis oils rich in aromatics, as well as more gas products (as compared with low severity pyrolysis). As will be appreciated by one of skill in the art, a high severity pyrolysis process will lead to the formation of more olefins and diolefins. Those olefins and diolefins cannot easily be recovered. The hydrotreatment of the present disclosure is therefore required.

A pyrolysis product can be recovered as an effluent from the pyrolysis unit and conveyed (e.g., flowed, for example via pumping, gravity, pressure differential, etc.) to a pyrolysis separating unit. The pyrolysis product can be separated in the pyrolysis separating unit into a pyrolysis gas stream and a pyrolysis plastic oil further used in step a) of the present disclosure. The pyrolysis separating unit may comprise any suitable gas-liquid separator, such as a vapor-liquid separator, oil-gas separators, gas-liquid separators, degassers, scrubbers, traps, flash drams, compressor suction drams, gravity separators, centrifugal separators, filter vane separators, mist eliminator pads, liquid-gas coalescers, distillation columns, and the like, or combinations thereof.

With regards to the steam cracker, it is known per se in the art. The feedstock of the steam cracker in addition to the stream obtained via the inventive process can be ethane, liquefied petroleum gas, naphtha or gasoils. Liquefied petroleum gas (LPG) consists essentially of propane and butanes. Gasoils have a boiling range from about 200 to 350° C., consisting of C10 to C22 hydrocarbons, including essentially linear and branched paraffins, cyclic paraffins and aromatics (including mono-, naphtho- and poly-aromatic).

In particular, the cracking products obtained at the exit of the steam cracker may include ethylene, propylene and benzene, and optionally hydrogen, toluene, xylenes, and 1,3-butadiene.

In a preferred embodiment, the outlet temperature of the steam cracker may range from 800 to 1200° C., preferably from 820 to 1100° C., more preferably from 830 to 950° C., more preferably from 840° C. to 920° C. The outlet temperature may influence the content of high value chemicals in the cracking products produced by the present process.

In a preferred embodiment, the residence time in the steam cracker, through the radiation section of the reactor where the temperature is between 650 and 1200° C., may range from 0.005 to 0.5 seconds, preferably from 0.01 to 0.4 seconds.

In a preferred embodiment, steam cracking is done in presence of steam in a ratio of 0.1 to 1.0 kg steam per kg of hydrocarbon feedstock, preferably from 0.25 to 0.7 kg steam per kg of hydrocarbon feedstock in the steam cracker, preferably in a ratio of 0.35 kg steam per kg of feedstock mixture, to obtain cracking products as defined above.

In a preferred embodiment, the reactor outlet pressure may range from 500 to 1500 mbars, preferably from 700 to 1000 mbars, more preferably may be approx. 850 mbars. The residence time of the feed in the reactor and the temperature are to be considered together. A lower operating pressure results in easier light olefins formation and reduced coke formation. The lowest pressure possible is accomplished by (i) maintaining the output pressure of the reactor as close as possible to atmospheric pressure at the suction of the cracked gas compressor (ii) reducing the pressure of the hydrocarbons by dilution with steam (which has a substantial influence on slowing down coke formation). The steam/feedstock ratio may be maintained at a level sufficient to limit coke formation.

Effluent from the steam cracker contains unreacted feedstock, desired olefins (mainly ethylene and propylene), hydrogen, methane, a mixture of C4's (primarily isobutylene and butadiene), pyrolysis gasoline (aromatics in the C6 to C8 range), ethane, propane, di-olefins (acetylene, methyl acetylene, propadiene), and heavier hydrocarbons that boil in the temperature range of fuel oil (pyrolysis fuel oil). This cracked gas is rapidly quenched to 338-510° C. to stop the pyrolysis reactions, minimize consecutive reactions and to recover the sensible heat in the gas by generating high-pressure steam in parallel transfer-line heat exchangers (TLE's). In gaseous feedstock-based plants, the TLE-quenched gas stream flows forward to a direct water quench tower, where the gas is cooled further with recirculating cold water. In liquid feedstock-based plants, a prefractionator precedes the water quench tower to condense and separate the fuel oil fraction from the cracked gas. In both types of plants, the major portions of the dilution steam and heavy gasoline in the cracked gas are condensed in the water quench tower at 35-40° C. The water-quench gas is subsequently compressed to about 25-35 Bars in 4 or 5 stages. Between compression stages, the condensed water and light gasoline are removed, and the cracked gas is washed with a caustic solution or with a regenerative amine solution, followed by a caustic solution, to remove acid gases (CO2, H2S and SO2). The compressed cracked gas is dried with a desiccant and cooled with propylene and ethylene refrigerants to cryogenic temperatures for the subsequent product fractionation: front-end demethanization, front-end depropanization or front-end deethanization.

The disclosure can be further defined using the following embodiments:

Embodiment 1. Process for the purification of a hydrocarbon stream comprising the following steps:a) Providing a hydrocarbon stream having a diene value of at least 1.5 g I2/100 g as measured according to UOP 326 and a bromine number of at least 5 g Br2/100 g as measured according to ASTM D1159 and containing at least 10 wt % of pyrolysis plastic oil the other part of said hydrocarbon stream being a first diluent;b) Optionally putting in contact the effluent obtained at the previous step with a silicon and/or metals and/or phosphorous and/or halogenates trap;c) Performing a first hydrotreating step at a temperature of at most 200° C.d) Optionally putting in contact the effluent obtained at the previous step with a silicon and/or metals and/or phosphorous and/or halogenates trap;e) performing a second hydrotreating step at a temperature of at least 200° C.,f) recovering a purified hydrocarbon stream.

Embodiment 2. Process according to previous embodiment wherein said pyrolysis plastic oil in said hydrocarbon stream has a starting boiling point of at least 15° C., and a final boiling point of preferably 560° C., more preferably 450° C. even more preferably 350° C., the most preferred 250° C., and/or said pyrolysis plastic oil has a diene value of at least 1.5, preferably 2, even more preferably 5 g I2/100 g, to at most 50 g I2/100 g as measured according to UOP 326, and/or contains more than 2 ppm wt of metals and/or said hydrocarbon stream contains preferably at least 25 wt %, even more preferably at least 50 wt %, even more preferably at least 75 wt % of said pyrolysis plastic oil and preferably at most 80 wt % of pyrolysis plastic oil, and/or at most 90 wt % preferably at most 95 wt %, even more preferably at most 100 wt % of said pyrolysis plastic oil.

Embodiment 3. Process according to any of the preceding embodiments wherein the weight concentration of said pyrolysis plastic oil in said hydrocarbon stream is chosen so that the total content of olefins, alkynes and diolefins in said hydrocarbon stream at the inlet of the second hydrotreatment is at most 20 wt %, preferably at most 15 wt %, most preferably at most 10 wt %.

Embodiment 4. Process according to any of the preceding embodiments wherein concerning said first hydrotreating step of said hydrocarbon stream one or more of the following statements is true:The inlet temperature ranges from 25 to 200° C.,The LHSV ranges from 1 to 10 h−1, preferably from 1 to 6 h−1, even more preferably from 2 to 4 h−1,The pressure ranges from 10 to 90 barg, preferably from 15-50 barg or preferably from 25 to 40 barg in presence of H2, and/or the molar ratio of H2 to the total molar sum of alkynes and dienes in said hydrocarbon stream is of at least 1.5, preferably at least 2, most preferably at least 3 to at most 15Said first hydrotreating step is performed in one or more catalyst bed with preferably an overall temperature increase of at most 150° C., more preferably of at most 100° C., and/or a temperature increase of at most 100° C., more preferably of at most 50° C. for each catalyst bed, with preferably intermediary quench between said catalyst beds, said quench being preferably performed with H2 or with said purified hydrocarbon stream recovered at step f)said first step is performed in a fixed bed reactor preferably over a catalyst that comprises at least one metal of group VIII, preferably selected from the group of Pt, Pd, Ni and/or mixture thereof on a support such as alumina, titania, silica, zirconia, magnesia, carbon; preferably said catalyst is a Ni based catalyst being a passivated after its reduction using preferably di-alkyl-sulfide such as DiMethylSulfide (DMS) or DiEthylSulfide (DES) or thiophenic compounds.said first step can also be performed in a fixed bed reactor preferably over a catalyst that comprises at least one metal of group VIB as for example Mo, W in combination or not with a promotor selected from at least one metal of group VIII and/VIIIB as for example Ni and/or Co, and/or mixture thereof, these metals being used in sulfided form and preferably supported on alumina, titania, zirconia, silica, carbon and/or mixtures thereofthe effluents obtained at the exit of said first hydrotreating step has a diene value of at most 1.5 g I2/100 g, preferably at most 1.0 g I2/100 g even more preferably at most 0.5 g I2/100 g.

Embodiment 5. Process according to any of the preceding embodiments wherein said trap of step b) is a silicon trap working at a temperature ranging from 20 to 100° C. and/or a LHSV between 1 to 10 h−1, and/or a pressure ranging from 1 to 90 barg and/or said trap of step d) is a silicon trap working at a temperature of at least 200° C., and/or a LHSV between 1 to 10 h−1, and/or a pressure ranging from 10 to 90 barg in presence of H2.

Embodiment 6. Process accord to any of the preceding embodiments wherein concerning the second hydrotreating step one or more of the following statements is true:No further hydrogenation step is necessary after said second hydrotreating step, preferably the concentration of olefins as measured via the bromine number in said purified hydrocarbon stream is at most 5.0, preferably at most 2.0 g Br2/100 g, more preferably at most 1.5 g Br2/100 g even more preferably at most 0.5 g Br2/100 g as measured according to ASTM D1159Said second hydrotreating step is performed in one or more catalyst bed with preferably an overall temperature increase of at most 100° C., and/or a temperature increase of at most 50° C. over each catalyst bed, with preferably intermediary quench between said catalyst beds, said quench being preferably performed with H2 or with said purified hydrocarbon stream recovered at step f)The inlet temperature is of at least 200° C. and at most 500° C.,The LHSV is between 1 to 10 h−1, preferably 2 to 4 h−1the pressure ranges from 10 to 90 barg in presence of H2Said second hydrotreating step is performed over a catalyst that comprises at least one metal of group VIB as for example Mo, W in combination or not with a promotor selected from at least one metal of group VIII and/or VIIIB as for example Ni and/or Co, and/or mixture thereof, preferably these metals being used in sulfided form and supported on alumina, titania, zirconia, silica, carbon and/or mixtures thereofthe ratio H2/hydrocarbon ranges from 200 NL/L to 900 NL/L, preferably in the presence of at least 0.005 wt %, preferably 0.05 wt % even more preferably 0.5 wt % of sulphur, being preferably H2S or organic sulfur compounds, in the stream; and/oron the top of the second hydrotreating step a silicon trap is present working at a temperature of at least 200° C., and/or a LHSV between 1 to 10 h−1, and/or a pressure ranging from 10 to 90 barg in presence of H2; optionally followed by a metal trap working at a temperature of at least 200° C., a LHSV between 1 to 10 h−1, a pressure ranging from 10 to 90 barg in presence of H2.

Embodiment 7. The process according to any of the preceding embodiments wherein said pyrolysis plastic oil and/or said hydrocarbon stream of step a) is treated before step b) in one or more of the followed pre-treatment unit:In a desalting unit to remove water-soluble saltsIn an impurities removal treatment step to remove silicon, phosphorous, metals and/or halogenated compounds, via preferably a solvent extraction or preferably in a guard bed, said guard bed preferably working at a temperature of at most 200° C., and/or a LHSV between 1 to 10 h−1, and/or a pressure ranging from 1 to 90 barg either in presence of H2 or in the absence of H2 and/or said guard bed is followed by a metal trap working at a temperature of at least 200° C., and/or a LHSV between 1 to 10 h−1, and/or a pressure ranging from 1 to 90 barg in presence of H2In a separation unit to extract the particles and gums by filtration, centrifugation or a combination of the two technics; and/orIn a dewatering unit to remove water in said hydrocarbon stream to reach a water content of less than 0.1% vol preferably of less than 0.05% vol according to ASTM D95.

Embodiment 8. The process according to any of the preceding embodiments wherein before performing the first and/or the second hydrotreating a further dilution is performed with the help of a diluent, said diluent being preferably a second hydrocarbon stream having a boiling range between 50° C. and 150° C. or a boiling range between 150° C. and 250° C. or a boiling range between 200° C. and 350° C. or with the effluent of said first and/or said second hydrotreating or any mixture thereof; said diluent being added to be at a concentration of at most 80 wt %, preferably at most 50 wt % and optionally said diluent is separated at the outlet of said first and/or of said second hydrotreating by a flash, or a distillation and preferably recycled at the inlet of said first and/or of said second hydrotreating and/or said diluent has preferably a bromine number of at most 5 g Br2/100 g, and/or a diene value of at most 0.5 g I2/100 g and/or a sulfur content of at most 1000 ppm wt.

Embodiment 9. The process according to any of the preceding embodiments wherein the stream entering the second hydrotreating is further diluted with any stream containing paraffins with an optional addition of a sulfur component, for instance DMDS, so that the concentration of sulfur is of at least 0.005 wt % of sulfur, preferably 0.05 wt % of sulfur to at most 0.5 wt %.

Embodiment 10. The process according to any the preceding embodiments wherein said purified hydrocarbon stream obtained at step f) is further mixed with naphtha, gasoil or crude oil to have a pyrolysis plastic oil concentration at the inlet of the steam cracker ranging from 0.01 wt % to at most 50 wt %; preferably 0.1 wt % to 25 wt % even more preferably 1 wt % to 20 wt % and sent to a steam cracker to produce olefins, such as ethylene and propylene, and aromatics.

Embodiment 11. The process according to the embodiments 1 to 9 wherein the purified hydrocarbon stream obtained at step f) is sent directly to the steam cracker without further dilution to produce olefins, such as ethylene and propylene, and aromatics.

Embodiment 12. The process according to any of the preceding embodiments wherein the part of effluent of the second hydrotreating step having an initial boiling point higher than 200° C., preferably higher than 300° C. even more preferably higher than 350° C. is sent to a FCC, or an hydrocracking unit, or a coker or a visbreaker or blended in crude oil or crude oil cut to be further refined.

Embodiment 13. The process according to any of the preceding embodiments wherein said pyrolysis plastic oil of step a) is originating from the stream of pyrolyzed waste plastic for which the C1 to C4 hydrocarbons have been removed and/or the components having a boiling point higher than 350° C. have been removed and/or preferably further converted into a FCC, or an hydrocracking unit, a coker or a visbreaker or blended in crude oil or crude oil cut to be further refined.

Embodiment 14. The process according to any of the preceding embodiments wherein the effluent obtained after said second hydrotreating step is further hydrocracked at a temperature of 350-430° C., a pressure of 30-180 barg, a LHSV of 0.5-4 h−1, and/or under a H2 to hydrocarbons ratio of 800-2000 NL:L to reduce the final boiling point with at least 10%.

Embodiment 15. The process according to any of the preceding embodiments wherein said first diluent is selected from a naphtha and/or a paraffinic solvent and/or a diesel or a straight run gasoil, containing at most 1 wt % of sulfur, preferably at most 0.1 wt % of sulfur, and/or an hydrocarbon stream having a boiling range between 50° C. and 150° C. or a boiling range between 150° C. and 250° C. or a boiling range between 200° C. and 350° C. having preferably a bromine number of at most 5 g Br2/100 g, and/or a diene value of at most 0.5 gI2/100 g even more preferably said first diluent is said purified hydrocarbon stream recovered at step e) or any combination thereof.

EXAMPLES

The embodiments of the present disclosure will be better understood by looking at the different examples below.

Example 1: 2ndStage Hydrotreating with Active Guard Bed to Trap Silicon Compound

The tests were performed with a sulfided NiMo catalyst by using a mixed feed C6/C7-C9 (60% wt-40% wt) doped with silicon compounds (˜10 wppm of hexamethylcyclotrisiloxane (HMCTS) and ˜10 wppm of octamethylcyclotetrasiloxane (OMCTS) representative, according to the literature1, of the decomposition of polysiloxanes. The table below gives more details about the feed used:

60% C6 + 40% C7-C9(Pygas unit Feed)doped with HMCTS and OMCTSDensity @15° C.g/ml0.7930Sulfur contentwppm<8Bromine NumbergBr2/100 g4.9C6 Oligomerwt %0.58Final Boiling Point° C.231Si contentHMCTS (ppm)11.3OMCTS (ppm)11.4UOP 744 methodNon aromatics(wt %)37.3Benzene(wt %)27.8Toluene(wt %)17.7Ethyl-benzene(wt %)0.7Xylenes o-m-p(wt %)1.9Other aromatics(wt %)10.2

The tests were performed at two temperatures 245° C. and 290° C. and two LHSV. The different operating conditions are summarized in Table 2.1Coker naphtha hydrotreating, R. Breivik and R. Egebjerg, Haldor TopsØe S/A, PTQ Q12008, 69-74.

PressureLHSVTH2/HCLiquid flowGas flowCondition(bars)(h−1)(° C.)(NI/I)(ml/h)(NI/h)Days127.71.124529011031.962290631031.9100031.93

The table here below shows the organic silicon content in the feed and in the effluents for the different conditions.

HMCTSOMCTSSi by XRFBr number/indexPollutants ppm11.3 ppm11.4 ppm9.8 ppm4.9 gBr2/100 gT (° C.)245<200 ppb.<200 ppb<1 ppm<50 mgBr2/100 g290LHSV = 1.1 h−1<200 ppb<200 ppb<1 ppm<50 mgBr2/100 gLHSV = 10 h−1<200 ppb<200 ppb<1 ppp<50 mgBr2/100 g

In terms of Si capture, the trap is convenient and efficient, even at higher LHSV.

Example 2: 2ndStage Hydrotreating with Active Guard Bed to Trap Silicon Compound and Hydrogenate

The tests were performed on a pyrolysis oil with a sulfided NiMo catalyst in the following conditions.

Pressure (barg)27.7LHSV (h−1)1.1Feed Flowrate (ml/h)55H2/HC (NI/l)290H2 flowrate (NI/h)16Temperature (° C.)Start Of Run: 245° C. The temperaturewas then gradually increased

The Pyrolysis oil was diluted in an inert product (Isoparaffinic cut) and characterized.

FeedFEED-50%IBP-FBP (° C.)83-427MV15 (g/ml)0.7806S by UV (ppm)6.2N (ppm)103.5Chlorine (ppm)44Metals par ICP-AES (ppm)Fe: 4, K: 1Si by XRF (ppm)35Br Number_(gBr/100 g)33Diene Value (gI2/100 g)1.7

The table here below show the results on this catalyst which is able to hydrogenate the double bonds and capture the silicon.

H2 = 16 NI/hBr NumberHC = 43 g/hSi_XRFT (° C.)(gBr2/100 g)Abattement(ppm)AbattementFEED3-10033 (±3)ER 10%35 (±3)ER 10%245° C.7.278%294%255° C.6.680%391%265° C.5.982%197%275° C.6.481%197%285° C.7.079%197%290° C.8.973%197%
On the effluent at 255° C., a molecular siloxane speciation was realized by GC-MS-SIM which highlight the abatements of siloxane molecules.

FEED-50%Effluent at 255° C.Si par XRF (ppm)353Spéciation siloxane(ppm)(ppm)Hexamethylcyclotrisiloxane16.5<Octamethylcyclotetrasiloxane10.5<Decamethylcyclopentasiloxane3.5<Dodecamethylcyclohexasiloxane2<Hexamethyldisiloxane<<Octamethyltrisiloxane<<Decamethyltetrasiloxane<<Dodecamethylpentasiloxane<<Somme (ppmSi)32.5<1
This example shows that a NiMo catalyst is able to both hydrogenate the olefins and at the same time trapping the siloxanes.

Example 3: Liquid Phase First Stage Hydrotreatment

The tests were performed using a pyrolysis plastic oil cut having a boiling point ranging from 70° C. to 460° C., a DV of about 4 g I2/100 g, a nitrogen content of about 210 wtppm and a sulfur content of about 20 ppm. A Ni on alumina catalyst was used in dilution 1:2 with silicon carbide 0.21 mm as diluent (50 ml of catalyst for 100 ml of SiC). The Nickel catalyst was dried under nitrogen (50 NI/h) at 180° C. and reduced under hydrogen (minimum 20 NI/h) at about 400° C. during min 15 h; then the temperature was reduced till 180° C. and the hydrogen was replaced by nitrogen to purge the reactor. Finally, the temperature was reduced to 50° C. and a paraffinic feed was injected to stabilize the catalyst.

The pyrolysis plastic oil cut was used pure.

The test was performed in the following operating conditions.

Pression (barg)30LHSV (h−1)2Q liquid (ml/h)100Q_H2 (NI/h)3 moles of hydrogen per mole of dienes.Temperature (° C.)Start Of Run: 50°

The temperature was increased till having the lowest DV in the liquid effluent. The Bromine number (BrN) is mentioned for information and to highlight that not all the olefins have been hydrogenated in these conditions.

DV(gI2/BrN(gBr2/Si by XRF100 g)100 g)(ppm)Feed4.1 ((±0.5)60.1 (±6)71 (±7)Tinlet (° C.)EffluentEffluentEffluent501.946.3—601.854.9—900.147.275

No noticeable exotherm was observed during the test, whatever the inlet Temperature considered. Increasing the temperature up to 120° C., allowed to decrease the Bromine number down to 42 g Br2/100 g. This example demonstrates that it is possible to hydrogenate the diolefins and the olefins of a pyrolysis plastic oil while maintaining the exothermicity in the catalyst bed at an acceptable level. The example demonstrates also that the silicon compound passes through this first hydrotreatment step.

Example 4 Liquid Phase First Stage Hydrotreatment

The tests were performed using a pyrolysis oil cut having a boiling point ranging from 20 to 250° C. A sulfided NiMo on alumina catalyst was used in dilution with silicon carbide at equal volumes.

The pyrolysis oil cut was diluted with a paraffinic diluent to have a MAV at the inlet of about 21 mg anhydride maleic/g (or a DV of about 5,4 gI2/100 g).

The test was performed in the following operating conditions.

Pression (barg)25LHSV (h−1)2liquide flow rate (ml/h)200H2/HC (NI/I)7H2 flow rate (NI/h)1.4Inlet Temperature (° C.)Start Of Run: 50°

No noticeable exotherm was observed during the test, whatever the inlet Temperature considered. The temperature was increased till having a MAV under 5.4 mg anhydride maleic/g (or a DV under 1,3 g I2/100 g) in the liquid effluent. This example demonstrates that it is possible to hydrogenate the diolefins and the olefins of a pyrolysis plastic oil while maintaining the exothermicity in the catalyst bed at an acceptable level.

Example 5: Adsorbents Used in Fixed Bed Reactor

It is foreseen that adsorbents will behave as it is presented in the results below. The tests were performed using a pyrolysis plastic oil cut having a boiling point ranging from 40° C. to 350° C. The water is expected to be below 100 ppm weight. The chlorine content of is expected to be in the range of about 200 ppm, the silicon content is expected to be in the range of about 100 ppm. The oxygen content of is expected in the range of about 1.0 wt %. The nitrogen is probably less than 2000 ppm wt. The adsorbent is chosen as being a promoted alumina (or active aluminium oxide) of spherical shape with 3.0 mm mean diameter with a surface area of 220 m2/g and a density of 0.75 kg/L. The adsorbent is disposed in a fixed bed under a continuous flow. Before the test the adsorbent shall be dried under nitrogen in up flow mode. The pyrolysis plastic oil shall be injected up flow. Dilution of the pyrolysis plastic oil with a first diluent can be done prior to the adsorption over the adsorbent. Alternatively, the pyrolysis plastic oil can be passed through the adsorbent without being diluted. This latter option was estimated in this example. The pyrolysis oil was injected in up flow mode at 20° C. under nitrogen blanketing.

Pyrolysis OilDensity @15° C.g/mL0.80Siliconppm100Oxygenwt %1.0Chlorineppm200Nitrogenppm2000

The tests were performed at ambient temperature (20° C.) and at three LHSV. The different operating conditions and performances expected are summarized in the following table, wherein weight percentage is given as the removed proportion of each measured element after treatment relative to the proportion of said element in the feedstock (here: plastic pyrolysis oil) before treatment. For sake of clarity, “100 wt %” means that the entirety of the component of interest has been removed:

PressureLHSVTOxygenChlorineSiliconNitrogenCondition(barg)(h−1)(° C.)wt %wt %wt %wt %110.5203015518211202812212312202310n.s.9*n.s. = not significant;

Impurities measurement was done at start of run. The overall oxygen uptake by the adsorbent is ranging from 2 to 15 wt % depending on operating conditions especially LHSV and the physical-chemical properties and nature of adsorbent used. This overall uptake corresponds to the maximal amount of oxygen containing impurities which can be trapped within the said adsorbent.

Very similar results can be obtained with silica gel having a spherical diameter of 5 mm, a surface area of about 500 m2/g, a density of 600 kg/m3 and a pore volume of about 0.42 cm3/g. The expected results with the same operating conditions are presented below

PressureLHSVTOxygenChlorineSiliconNitrogenCondition(barg)(h−1)(° C.)wt %wt %wt %wt %110.52020127302112015952231220117n.s.416*n.s. = not significant;

It appears from the examples described above that the promoted alumina and silica gel should allow to trap oxygen, chlorine, nitrogen and also silicon to a certain extend too.