The present invention relates to a process for purifying a pyrolysis oil comprising a hydrogenation step prior to a dehalogenation step, a production unit for carrying out said process and a purified pyrolysis oil obtained or obtainable by said process.

Currently, plastic waste is still largely landfilled or incinerated for heat generation. Chemical recycling is an attractive way to convert waste plastic material into useful chemicals. An important technique for chemically recycling plastic waste is pyrolysis. The pyrolysis is a thermal degradation of plastic waste in an inert atmosphere and yields value added products such as pyrolysis gas, liquid pyrolysis oil and char (residue), wherein pyrolysis oil is the major product. The pyrolysis gas and char can be used as fuel for generating heat, e.g. for reactor heating purposes. The pyrolysis oil can be used as source for syngas production and/or processed into chemical feedstock such as ethylene, propylene, C4 cuts, etc. for example in a (steam) cracker.

Typically, the plastic waste is mixed plastic waste composed of different types of polymers. The polymers are often composed of carbon and hydrogen in combination with other elements such as chlorine, bromine, fluorine, sulfur, oxygen and nitrogen that complicate recycling efforts. The elements other than carbon and hydrogen may be harmful during the further processing of the crude pyrolysis oil, since they may deactivate or poison catalysts used in the further processing of the pyrolysis oil. During (steam) cracking, halogen-containing compounds can damage the cracker by corrosion in that they release hydrogen halide. Sulfur-containing compounds can deactivate or poison catalysts used in the cracker, or can contaminate the cracker products. Nitrogen containing impurities may also poison downstream catalysts. In addition, they may cause a safety problem by forming explosive NOx when heated. When mixed plastics containing polyvinyl chloride (PVC) is thermally degraded, compounds having double carbon bonds and hydrogen chloride is formed. The hydrogen chloride liberated from PVC attacks the compounds having carbon-carbon double bonds leading to the formation of chlorinated organic compounds. Plastic waste typically contains heteroatom containing additives such as stabilizers and plasticizers that have been incorporated to improve the performance of the polymers. Such additives also often comprise nitrogen, halogen and sulfur containing compounds and heavy metals. For example, waste engine oils, transformer oils, hydraulic oils and machine oils may contain heavy metal abrasion. The heavy metals are often toxic and the quality of the pyrolysis oil is reduced by the presence of heavy metal impurities. Furthermore, plastic waste often may be uncleaned plastics with residue that may also contain elements other carbon and hydrogen. Therefore, the reduction of the nitrogen, sulfur, halogen content in the pyrolysis oil as well as the heavy-metal content is essential for any profit-generating processing of the pyrolysis oil. Especially, a high quality pyrolysis oil rich in carbon and hydrogen and low in elements other than carbon and hydrogen is preferred as feedstock to prevent catalyst deactivation and corrosion problems in downstream refinery processes.

WO 2017/083018 A1 discloses a process for reducing chloride content of a hydrocarbon feed stream. Further, FR 3 103 822 discloses a process for treating pyrolysis oil for subsequent steam cracking, said process comprises a hydrogenation step prior to a hydroprocessing step. However, there is still a need to provide improved process for purifying pyrolysis oil obtained from plastic waste.

Therefore, there is a need to provide a process for purifying pyrolysis oils, preferably obtained from plastic waste, in particular by reducing the diene content as well as chlorine, nitrogen and sulfur contents. Indeed, there is a need to provide high value purified pyrolysis oils while using an economic process.

It was surprisingly found that the process of the present invention permits to reduce the diene content as well as chlorine, nitrogen and sulfur contents, such reduced amounts being particularly adapted for subsequent steam cracking. Further, it was surprisingly found that the process of the present invention was an economic process providing high value purified pyrolysis oils.

Further, it was found that dehalogenation of pyrolysis oil may turn out to be an essential step to sustain steady reactor downstream operation. A reason for the instability is believed to be the formation of ammonium halide such as ammonium chloride which may desublimate upon lowering the temperature of the effluent of a hydroprocessing step. As nitrogen is typically present in larger amounts in the pyrolysis oil feed than halogene such as chlorine, and can only be removed via hydroprocessing and then is present as ammonia, it is believed to be essential to lower the halogen content upstream of a hydroprocessing step, thus avoiding desublimation of ammonium halide and and reducing its absolute content.

Furthermore, it is believed that corrosive attack of technical equipment can be reduced as e.g. HCI is produced from organic chlorides under hydroprocessing conditions. Ammonium chloride itself may cause corrosion, too.

Yet further, it was found that at operations as dechlorination and/or hydroprocessing, catalytic and/or adsorbent performance suffers from increased contents of dienes and/or conjugated aromatic olefins. Their oligomerization respective polymerization is believed to be caused by increased temperature and/or radical forming compounds as e.g. peroxides or other impurities as acids or metal salts acting such as lewis acids. It was found that such oligo- or polymerization can coke the catalyst and/or the adsorbent or block their pore system reducing accessibility of reactants to the catalyst and/or adsorbent.

Therefore, the present invention relates to a process for purifying a pyrolysis oil, the process comprising:

(i) providing a stream SO comprising a pyrolysis oil, the pyrolysis oil comprising one or more halogenated organic compounds and one or more organic compounds comprising conjugated double bonds;

(ii) subjecting the stream SO provided in (i) to hydrogenation in at least one reaction zone Z1 containing a heterogeneous hydrogenation catalyst, obtaining a stream S1 being deplet- ed, compared to SO, in the one or more organic compounds comprising conjugated double bonds;

(iii) subjecting the stream S1 obtained from (ii) to dehalogenation in at least one dehalogenation zone Z2 downstream of Z1, obtaining a stream S2 being depleted, compared to S1 , in the one or more halogenated organic compounds.

The term “dehalogenation” as used in the context of the present invention generally comprises “dechlorination”, “debromination” as well as “defluorination”. According to the present invention, the term “dehalogenation” preferably comprises “dechlorination”. If, e.g., the pyrolysis oil to be subjected to the process according to the present invention does not contain brominated organic compounds and fluorinated organic compounds, but only chlorinated compounds as halogenated organic compounds, the term “dehalogenation” would be directed to “dechlorination”, and the process of the invention would a process for purifying a pyrolysis oil, the process comprising:

(i) providing a stream SO comprising a pyrolysis oil, the pyrolysis oil comprising one or more chlorinated organic compounds and one or more organic compounds comprising conjugated double bonds;

(ii) subjecting the stream SO provided in (i) to hydrogenation in at least one reaction zone Z1 containing a heterogeneous hydrogenation catalyst, obtaining a stream S1 being depleted, compared to SO, in the one or more organic compounds comprising conjugated double bonds;

(iii) subjecting the stream S1 obtained from (ii) to dechlorination in at least one dechlorination zone Z2 downstream of Z1 , obtaining a stream S2 being depleted, compared to S1 , in the one or more chlorinated organic compounds; wherein a preferred step (iii) would comprise

(111.1) introducing a gas stream G1 into Z2 preferably being an adsorption zone, more preferably a gas stream comprising one or more of hydrogen and nitrogen, more preferably hydrogen;

(111.2) introducing the stream S1 obtained from (ii) into Z2;

(111.3) bringing S1 in contact with G1 and a heterogeneous adsorbent material, suitable for adsorbing chloride comprised in at least one of the one or more chlorinated organic compounds, comprised in Z2, obtaining a stream S2 being depleted, compared to S1 , in the one or more chlorinated organic compounds;

(111.4) removing S2 from Z2.

Generally, the pyrolysis oil comprised in SO can be a non-treated crude pyrolysis oil. Further, the pyrolysis oil comprised in SO can be a suitably pretreated crude pyrolysis oil. Such suitable pretreatment procedures are non-hydrogenation methods and include, but are not restricted to, distillation, dilution, precipitation, filtration, and extraction. The said crude pyrolysis oil can be subjected to one suitable pretreatment procedure, or to two or more suitable pretreatment procedures. If the pretreatment is, for example, a distillation, it is possible to enrich the pyrolysis oil with respect to at least one of the one or more organic compounds comprising conjugated double bonds, or with respect to at least one of the one or more halogenated organic compounds, or with respect to at least one of the one or more organic compounds comprising conjugated double bonds and at least one of the one or more halogenated organic compounds. If the pretreatment is, for example, a dilution, it is possible to add one or more alkanes. By doing so, it may be possible to precipitate one or more asphaltenes from the mixture resulting from dilution which may then be removed by filtration. If the pretreatment is, for example, an extraction, it is possible to use an aqueous extraction medium, such as an acidic aqueous extraction medium or a basic extraction medium.

Generally, from 1 to 100 weight-% or from 5 to 100 weight-% or from 10 to 100 weight-% or from 20 to 100 weight-% or from 30 to 100 weight-% or from 40 to 100 weight-% or from 50 to 100 weight-% or from 60 to 100 weight-% or from 70 to 100 weight-% or from 80 to 100 weight- % or from 90 to 100 weight-% of SO may consist of pyrolysis oil.

Preferably from 95 to 100 weight-%, more preferably from 98 to 100 weight-%, more preferably from 99 to 100 weight-%, of SO consist of pyrolysis oil.

It is possible that from 99.5 to 100 weight-% or from 99.8 to 100 weight-% or from 99.9 weight- % of SO consist of pyrolysis oil.

Preferably the pyrolysis oil according to (i) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of from 0.1 to 75 g(l2)/100 g, more preferably from 0.4 to 60 g(l2)/100 g, more preferably from 1 to 30 g(l2)/100 g of the pyrolysis oil, determined as described in Reference Example 1.

Preferably the one or more organic compounds comprising conjugated double bonds comprise one or more organic compounds according to formula (I)

R ¹ R ² C ¹ =C ² R ³ -C ³ R ⁴ =X (I) wherein =X is =0, =S, =NR ⁵ , or =C ⁴ R ⁶ R ⁷ , preferably =C ⁴ R ⁶ R ⁷ .

Preferably R ¹ , R ² , R ³ , R ⁴ , R ⁵ are, independently of each other, H, alkyl having from 1 to 6 carbon atoms, alkenyl having from 1 to 6 carbon atoms, or aryl having from 5 to 10 carbon atoms, more preferably H.

Preferably R ⁶ and R ⁷ are, independently of each other, H, alkyl having from 1 to 6 carbon atoms, alkenyl having from 1 to 6 carbon atoms, or aryl having from 5 to 10 carbon atoms, more preferably H. Alternatively, preferably either R ⁴ and R ⁶ or R ⁴ and R ⁷ are linked together, thus forming, together with C ³ =C ⁴ , an aromatic ring more preferably having 5 or 6 members.

Preferably the one or more organic compounds according to formula (I) comprise one or more of butadiene, isoprene, dienes having 5 or 6 carbon atoms, styrene, methyl styrene, indene, substituted styrene, substituted indene, and 3-methyl-2-butenal, more preferably comprise one or more of butadiene, isoprene, dienes having 5 or 6 carbon atoms, styrene, methyl styrene, indene, and 3-methyl-2-butenal.

As to the one or more organic compounds according to formula (I), it is preferred that they comprise styrene.

Preferably the pyrolysis oil according to (i) has a styrene content in the range of from 0.2 to 30 Area%, more preferably in the range of from 1 to 20 Area%, determined as described in Reference Example 8.

Any halogenated organic compounds can be comprised in the pyrolysis oil comprised in SO provided in (i). Preferably the one or more halogenated organic compounds comprise one or more of mono-, oligo- or polyhalogenated aromatic compounds, alkyl halides and alkenyl halides.

In the context of the present invention, the pyrolysis oil to be purified can have any chlorine content. It is however preferred that the pyrolysis oil has a total chlorine content in the range of from 30 to 3,000 wppm (ppm by weight), more preferably from 30 to 500 wppm, more preferably from 30 to 200 wppm, determined as described in Reference Example 2.1 .

In the context of the present invention, the pyrolysis oil to be purified can have any chloride content. It is however preferred that the pyrolysis oil has a chloride content of at most 100 wppm, more preferably in the range of from 0 to 30 wppm, determined as described in Reference Example 2.2.

In the context of the present invention, the pyrolysis oil to be purified can have any nitrogen content. It is however preferred that the pyrolysis oil has a nitrogen content in the range of from 50 to 20,000 wppm (ppm by weight), more preferably from 50 to 5,000 wppm, more preferably from 100 to 4,000 wppm, determined as described in Reference Example 3.

In the context of the present invention, the pyrolysis oil to be purified can have any sulfur content. It is however preferred that the pyrolysis oil has a sulfur content in the range of from 50 to 30,000 ppm by weight (wppm), preferably from 50 to 5,000 wppm, more preferably from 100 to 3,000 wppm, determined as described in Reference Example 4.

In the context of the present invention, the pyrolysis oil is preferably obtained from plastic waste. In the context of the present invention, the “plastic waste” to be pyrolyzed typically is mixed or pre-sorted plastic waste. However, it is also possible to use plastic waste resulting from tires, plastic waste which is pure polymeric plastic waste, or film waste, including soiling, adhesive materials, fillers, residues, etc. The pyrolysis oil to be purified can typically comprise a solid phase and a liquid phase, wherein the liquid phase includes an organic phase and an aqueous phase. For example, a weight ratio between the aqueous phase and the organic phase in the liquid phase of the pyrolysis oil can be in the range of from 0.01 :1 to 3.2:1 , preferably in the range of from 0.05:1 to 3:1.o

In the context of the present invention, the stream SO is preferably a liquid stream. It is conceivable that the pyrolysis oil is at least partially in the form of a wax which, prior to be subjected to the process of the present invention, is suitably liquefied.

Pre-Hydrogenation

As to (ii), the stream SO subjected to hydrogenation in (ii) has preferably a temperature in the range of from 60 to 250 °C, more preferably from 80 to 220 °C, more preferably from 100 to 200 °C.

In the context of the present invention, any heterogeneous hydrogenation catalyst can be used as far as it permits to hydrogenate (pre-hydrogenation) according to (ii) the stream SO comprising the pyrolysis oil to be purified prior to the dehalogenation according to (iii). Preferably, the heterogeneous hydrogenation catalyst according to (ii) comprises an element of the groups 8 to 12, more preferably 8 to 10, more preferably 9 and 10, of the periodic table of elements, more preferably an element selected from the group consisting of Ni, Pd and Co, more preferably from the group consisting of Ni and Pd.

Preferably, the heterogeneous hydrogenation catalyst according to (ii) further comprises a support material for said element of the groups 8 to 12 of the periodic table of elements, wherein the support material is preferably selected from the group consisting of an oxidic material and carbon, wherein the oxidic material is more preferably one or more of alumina, silica, magnesia, zirconia, titania, a zeolitic material, a silica-alumina phosphate (SAPO) material, zinc oxide, sodium oxide, mixed silica-alumina, zeolite and calcium oxide, more preferably alumina.

As to the heterogeneous hydrogenation catalyst according to (ii), it is preferred that said catalyst comprises Ni, more preferably in an amount, calculated as NiO, in the range of from 0.5 to 70 weight-%, more preferably from 0.75 to 45 weight-%, more preferably from 1 to 20 weight-%, based on the total weight of the hydrogenation catalyst.

Preferably, the heterogeneous hydrogenation catalyst according to (ii) further comprises an element of the group 6 of the periodic table of elements, wherein the element of the group 6 is more preferably one or more of Mo and W, more preferably Mo. Preferably, the heterogeneous hydrogenation catalyst according to (ii) comprises from 1 to 40 weight-%, more preferably from 2 to 35 weight-%, more preferably from 3 to 30 weight-% of an oxide of said element of the group 6, more preferably Mo oxide or W oxide, based on the total weight of the hydrogenation catalyst.

Preferably, the heterogeneous hydrogenation catalyst according to (ii) comprises Ni and Mo supported on a support material, more preferably a support material as defined in in the foregoing. More preferably, the hydrogenation catalyst comprises Ni and Mo supported on alumina.

As to the heterogeneous hydrogenation catalyst according to (ii), it is alternatively preferred that said catalyst comprises Pd, more preferably in an amount, calculated as elemental Pd, in the range of from 0.01 to 5 weight-%, more preferably from 0.1 to 1 weight-%, more preferably from 0.15 to 0.8 weight-%, based on the total weight of the catalyst.

Preferably, the heterogeneous hydrogenation catalyst according to (ii) further comprises a promoter, the promoter more preferably being one or more of an element of the groups 10 and 11 of the periodic table of elements, preferably one or more of Cu, Au, Ag, and Pt, more preferably one or more of Ag and Pt, more preferably Ag.

Preferably, the atomic ratio of the element of groups 8 to 12 of the periodic table, more preferably Pd, relative to the promoter is in the range of from 0.1 :1 to 10:1 , more preferably from 2:1 to 7:1 , more preferably from 2.5:1 to 6:1.

Preferably, the heterogeneous hydrogenation catalyst according to (ii) comprises Pd supported on a support material, preferably a support material as defined in the foregoing, wherein the support material is more preferably alumina or carbon, more preferably alumina.

In the context of the present invention, the heterogeneous hydrogenation catalyst according to (ii) preferably is in the form of extrudates, pellets, rings, spherical particles or spheres, more preferably spherical particles or extrudates.

Preferably, (ii) comprises

(11.1 ) introducing a gas stream GO into Z1 , the gas stream comprising H2;

(11.2) introducing the stream SO into Z1 ;

(11.3) bringing SO in contact with GO and the heterogeneous hydrogenation catalyst comprised in Z1 , obtaining a stream S1 being depleted, compared to SO, in the one or more organic compounds comprising conjugated double bonds;

(11.4) removing S1 from Z1.

Preferably, the gas stream GO has a temperature in the range of 100 to 250 °C, more preferably from 120 to 220 °C, more preferably from 140 to 200 °C. Preferably, the gas stream GO is introduced at a pressure in the range of from 10 to 100 bar(abs), more preferably from 15 to 90 bar(abs), more preferably from 20 to 80 bar(abs), more preferably in the range of from 20 to 55 bar(abs).

Preferably, from 70 to 100 weight-%, more preferably from 80 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream GO consists of H2. Further conceivable ranges are from 92 to 100 weight-% or from 94 to 100 weight-% or from 96 to 100 weight-% or from 98 to 100 weight-%.

According to (ii.1), GO is preferably introduced continuously or semi-continuously, more preferably continuously into Z1 .

According to (ii.2), SO is preferably introduced semi-continuously or continuously, more preferably continuously, into Z1.

Preferably, GO is introduced into Z1 according to (ii.1) for a period At prior to introducing SO into Z1 according to (ii.2).

In particular in case the hydrogenation catalyst comprises Ni, it may preferred to sulfidize the catalyst, preferably after drying, before it is employed as hydrogenation catalyst. Drying may be accomplished by subjecting the catalyst to a gas atmosphere, preferably containing nitrogen such as technical nitrogen, the gas atmosphere preferably having a temperature in the range of from 150 to 250 °C such as from 170 to 230 °C or from 190 to 210 °C. If drying is carried out semi-continuously or continuously, the gas hourly space velocity may be in the range of from 1 ,000 to 3,000 IT ¹ such as from 1 ,500 to 2,500 IT ¹ or from 1 ,800 to 2,200 IT ¹ . After drying, H2S is dosed preferably until saturation with sulfur is achieved, preferably at a temperature in the range of from 300 to 400 °C such as from 325 to 375 °C.

Alternatively, a hydrogenation catalyst comprising Ni can be presulfidized according to the following (or a similar) method: the catalyst is dried at 200 °C (temperature increase rate 1 K/min) for at least 2 h under flow of nitrogen(GHSV=2000/h) at atmospheric pressure until no water is condensed anymore downstream. Afterwards the catalyst is cooled down to 135 °C and hydrogen is fed to the reactor at a GHSV=2000/h and the reactor is pressurized (1 bar/min). After 1 h hexadecane spiked with 2 weight-% dimethyldisulfide is dosed with LHSV=2/h for 1 h. Afterwards temperature is increased with 0.25K/min to 350 °C and kept for 2h. Afterwards, temperature and pressure are adjusted before hexadecane solution dosing is stopped and feed is dosed.

If the hydrogenation catalyst comprises Pd, it may be preferred that it is activated under flow of hydrogen (GHSV=1000/h) at 50 to 130 °C (1 K/min), for example for 6 to 24 h such as 12 h, preferably at atmospheric condition. Upon catalyst reduction in larger reactor, hydrogen can be diluted by nitrogen to avoid excess temperature. During At, GO is preferably brought in contact with the heterogeneous hydrogenation catalyst comprised in Z1 , wherein GO has a temperature in the range of 50 to 250 °C, more preferably from 120 to 220 °C, more preferably from 140 to 200 °C.

In Z1 , the liquid hourly space velocity (LHSV) is preferably in the range of from 0.2 to 10 m ³ /(m ³ h), more preferably in the range of from 0.3 to 5 m ³ /(m ³ h), more preferably in the range of from 0.5 to 2 m ³ /(m ³ h), wherein the LHSV is defined as the volume flow of SO through Z1 (in m ³ /h) per volume of heterogeneous hydrogenation catalyst comprised in Z1 (in m ³ ).

Generally, according to the present invention, it is possible to recycle a portion of S1 , S1 ’, obtained from (ii) and removed from Z1 , back to Z1 as part of the starting material subjected to hydrogenation. In this case, not only the stream SO is subjected to hydrogenation, but also the additional stream ST. As far as ST is concerned, its LHSV in Z1 , relative to the LHSV of SO, LHSV(S1 ’):LHSV(S0), corresponding to the recycle ratio, is preferably in the range of from 1 :1 to 20:1 , such as from 1 :1 to 5:1 or from 5:1 to 10:1 or from 10:1 to 15:1 or from 15:1 to 20:1.

Generally, ST can be admixed with SO at every suitable position in the process, preferably upstream of Z1 . If a portion of S1 is recycled, the remaining portion of S1 is referred to herein as the stream S1 which is subjected to dehalogenation according to (iii).

Generally, it is preferred that the process further comprises

(ii .5) removing a gas stream GO’ from Z1 , the gas stream GO’ comprising H2.

Generally, according to the present invention, it is possible to recycle a portion of GO’, GO”, obtained from (ii) and removed from Z1 , back to Z1 as part of the starting materials. In this case, not only the stream GO’ is introduced into Z1 , but also the additional stream GO”. As far as GO” is concerned, its hourly space velocity, GHSV, in Z1 , relative to the GHSV of GO, LHSV(G0”):LHSV(G0), corresponding to the recycle ratio, is preferably in the range of from 1 :1 to 20:1 , such as from 1 :1 to 5:1 or from 5:1 to 10:1 or from 10:1 to 15:1 or from 15:1 to 20:1.

Preferably, the reaction zone Z1 is comprised in a continuous stirred tank reactor (CSTR) or a fixed bed reactor, more preferably in a fixed bed reactor, wherein the fixed bed reactor is more preferably a trickle bed reactor.

According to (ii), two or more reaction zones Z1 are preferably employed which are arranged serially and/or in parallel, wherein more preferably, one single reaction zone Z1 is employed according to (ii).

Preferably, the stream S1 obtained from (ii), and subjected to dehalogenation in (iii), comprises a reduced amount of the one or more organic compounds comprising conjugated double bonds of from 50 to 100%, more preferably from 70 to 100 %, more preferably from 75 to 100 %, compared to SO. The total amount of the one or more organic compounds comprising conjugated double bonds is determined as described in Reference Example 1. Preferably, the stream S1 obtained from (ii) and subjected to dehalogenation in (iii) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of from 0 to 3 g(l2)/100 g, preferably from 0 to 2 g(l2)/100g, more preferably from 0 to 1 g(l2)/100 g, more preferably from 0 to 0.25 g(l2)/100 g, more preferably from 0 to 0.1 g(l2)/100 g of the stream S1 , determined as described in Reference Example 1 .

Preferably, the stream S1 obtained from (ii), and subjected to dehalogenation in (iii), comprises a reduced styrene amount from 50 to 100%, more preferably from 70 to 100 %, more preferably from 75 to 100 %, compared to SO. The styrene amount is determined as described in Reference Example 8.

Preferably, the stream S1 obtained from (ii) and subjected to dehalogenation in (iii) has a styrene content in the range of from 0 to 1 .5 Area%, more preferably in the range of from 0 to 0.1 Area%, determined as described in Reference Example 8.

Dehalogenation

Preferably, the stream S1 subjected to dehalogenation in (iii) has a temperature in the range of from 150 to 450°C, more preferably from 200 to 400°C, more preferably from 250 to 350 °C.

Adsorption zone Z2

Preferably, the dehalogenation zone Z2 according to (iii) comprises, more preferably consists of an adsorption zone, more preferably comprising a heterogeneous adsorbent material suitable for adsorbing halide comprised in at least one of the one or more halogenated organic compounds, preferably comprised all of the one or more halogenated organic compounds.

In the context of the present invention, any heterogeneous adsorbent material can be used according to (iii) as far as it permits to adsorb halide comprised in at least one of the one or more halogenated organic compounds. Preferably the heterogeneous adsorbent material according to (iii) comprises one or more of a carbon-containing adsorbent material and an aluminum- containing adsorbent material, more preferably an aluminum-containing adsorbent material.

The carbon-containing adsorbent material is preferably a carbon-containing molecular sieve, more preferably activated charcoal. The aluminum-containing adsorbent material is preferably an alumina, an aluminum-containing molecular sieve, a silicoaluminophosphate, a silica- alumina hydrate or a hydrotalcite. The aluminum-containing molecular sieve is preferably an aluminosilicate, more preferably having a molar ratio of Si: Al, calculated as SiC^AhOs, in the range of from 2:1 to 10:1 , more preferably from 2:1 to 4:1. The silica-alumina hydrate preferably has weight ratio AhO3:SiO2 in the range of from 1 :1 to 10:1 , more preferably from 1 :1 to 2:1. The hydrotalcite is preferably an aluminum and magnesium containing hydrotalcite, more preferably an aluminum-magnesium hydroxycarbonate, preferably having a MgOAhOs weight ratio in the range of from 63:37 to 70:30.

Preferably, the heterogeneous adsorbent material comprises an hydrotalcite as defined in the foregoing.

It is also conceivable that the heterogeneous adsorbent material according to (iii) comprises an element of the groups 1 , 2, 11 and 12.

Preferably, the aluminosilicate mentioned above contains one or more of potassium oxide, sodium oxide, magnesium oxide and calcium oxide.

Preferably, according to the present invention, from 0 to 0.001 weight-%, preferably from 0 to 0.0001 weight-%, more preferably from 0 to 0.00001 weight-%, of the heterogeneous adsorbent material according to (iii) consist of Ni.

Preferably, the heterogeneous adsorbent material according to (iii) comprises particles characterized by a particle size distribution having a D50 value in the range of from 1 to 6,500 micrometers, more preferably from 2 to 2,000 micrometers, more preferably from 8 to 500 micrometers, more preferably from 10 to 50 micrometers or from 3 to 9 micrometers, the D50 particle size being determined as described in Reference Example 5.

More preferably, the heterogeneous adsorbent material being an hydrotalcite defined in the foregoing comprises particles characterized by a particle size distribution having a D50 value in the range of from 3 to 9 micrometers, the D50 particle size being determined as described in Reference Example 5.

In the context of the present invention, the heterogeneous adsorbent material according to (iii) has preferably an average pore volume in the range of from 0.1 to 5 ml/g, more preferably in the range of from 0.15 to 2 ml/g, the average pore volume being determined as described in Reference Example 6.

Preferably, the heterogeneous adsorbent material according to (iii) has a BET specific surface area in the range of from 50 to 1 ,000 m ² /g, more preferably in the range of from 100 to 900 m ² /g, more preferably in the range of from 150 to 600 m ² /g, the BET specific surface area being determined as described in reference Example 7.

According to the present invention, the adsorbent material can be suitably regenerated, if so desired.

Preferably, (ill) comprises (111.1 ) introducing a gas stream G1 into Z2 preferably being an adsorption zone, more preferably a gas stream comprising one or more of hydrogen and nitrogen, more preferably hydrogen;

(111.2) introducing the stream S1 obtained from (ii) into Z2;

(111.3) bringing S1 in contact with G1 and a heterogeneous adsorbent material, suitable for adsorbing halide comprised in at least one of the one or more halogenated organic compounds, comprised in Z2, obtaining a stream S2 being depleted, compared to S1 , in the one or more halogenated organic compounds;

(111.4) removing S2 from Z2.

Preferably, the gas stream G1 has a temperature in the range of 250 to 500°C, more preferably in the range of from 300 to 400 °C.

Preferably, the gas stream G1 is introduced at a pressure in the range of from 1 to 100 bar(abs), more preferably in the range of from 5 to 80 bar(abs), more preferably in the range of from 10 to 50 bar(abs).

Preferably, in Z2, the liquid hourly space velocity (LHSV) is in the range of from 0.2 to 10 m ³ /(m ³ h), more preferably in the range of from 0.3 to 5 m ³ /(m ³ h), more preferably in the range of from 0.5 to 2 m ³ /(m ³ h), wherein the LHSV is defined as the volume flow of S1 through Z2 (in m ³ /h) per volume of adsorbent material comprised in Z2 (in m ³ ).

Preferably, from 90 to 100 weight-%, more preferably from 95 to 100 weight-%, more preferably from 98 to 100 weight-%, of the gas stream G1 consists of H2. Alternatively, preferably from 90 to 100 weight-%, more preferably from 95 to 100 weight-%, more preferably from 98 to 100 weight-%, of the gas stream G1 consists of nitrogen.

According to (iii.1 ), G1 is preferably introduced continuously or semi-continuously, more preferably continuously, into Z2 and wherein according to (iii.2) S1 is introduced continuously or semi- continuously, more preferably continuously, into Z2.

Preferably, the adsorption zone Z2 is comprised in a continuous stirred tank reactor (CSTR) or a fixed bed reactor, more preferably in a fixed bed reactor, more preferably a trickle bed reactor, the reactor preferably comprising an adsorption bed comprising the heterogeneous adsorbent material.

According to (iii), two or more reaction zones Z2 are preferably employed which are arranged serially and/or in parallel or one single reaction zone Z2 is preferably employed according to (iii), more preferably one single reaction zone Z2 is employed according to (iii).

Preferably, the stream S2 obtained from (iii) has a reduced total chlorine content of from 50 to 100 %, preferably 60 to 100 %, more preferably from 75 to 100%, compared to SO and S1 . Preferably, the stream S2 obtained from (iii) has a total chlorine content in the range of from 0 to 200 wppm (ppm by weight), more preferably from 0 to 160 wppm, more preferably from 0 to 130 wppm, determined as described in Reference Example 2.1.

Preferably, the stream S2 obtained from (iii) has a chloride content of at most 40 wppm (ppm by weight), more preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 1 wppm, determined as described in Reference Example 2.2.

Preferably, the stream S2 obtained from (iii) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of 0 to 3 g(l2)/100 g, more preferably from 0 to 2 g(l2)/100 g, more preferably from 0 to 1 g(l2)/100 g, more preferably from 0 to 0.25 g(l2)/100 g, more preferably from 0 to 0.1 g(l2)/100 g of the stream S2, determined as described in Reference Example 1 .

Preferably, the stream S2 obtained from (iii) has a nitrogen content in the range of from 50 to 20,000 ppm by weight wppm, more preferably from 50 to 5,000 wppm, more preferably from 100 to 4,000 wppm, determined as described in Reference Example 3.

Preferably, the stream S2 obtained from (iii) has a sulfur content in the range of from 50 to 30,000 ppm by weight (wppm), more preferably from 50 to 5,000 wppm, more preferably from 100 to 3,000 wppm, determined as described in Reference Example 4.

Preferably, the stream S2 has a chlorine content, a nitrogen content, a sulfur content and a total amount of the one or more organic compounds comprising conjugated double bonds as defined in the foregoing.

Catalytic zone Z2

Alternatively, the dehalogenation zone Z2 according to (iii) comprises, more preferably is a catalytic zone, the catalytic zone more preferably comprising a heterogeneous dehalogenation catalyst, said catalyst preferably comprising one or more catalytically active elements of groups 8 to 12 of the periodic system of elements.

In this case, (iii) preferably comprises

(iii.1 ’) introducing a gas stream G1 into Z2 being a catalytic zone, more preferably introducing a gas stream comprising hydrogen and nitrogen, more preferably hydrogen;

(iii.2’) introducing the stream S1 obtained from (ii) into Z2;

(iii.3’) bringing S1 in contact with G1 and a heterogeneous dehalogenation catalyst comprised in Z2, obtaining a stream S2 being depleted, compared to S1 , in the one or more halogenated organic compounds;

(iii.4’) removing S2 from Z2;

(iii.5’) more preferably cooling S2 removed according to (iii .4’). Preferably, the stream S2 obtained from (iii), prior to being subjected to hydroprocessing according to (iv) as defined in the following, is subjected to extraction, more preferably using an aqueous extraction medium, obtaining a stream S2 being depleted in one or more dissolved halides comprised in S2 obtained from the dehalogenation zone Z2, said halides more preferably comprising one or more halides of N-containing organic compounds.

Hydroprocessing

The process of the present invention preferably further comprises, after (iii),

(iv) subjecting the stream S2 obtained from (iii) to hydroprocessing in at least one reaction zone Z3 downstream of Z2, Z3 comprising a heterogeneous hydroprocessing catalyst; obtaining a stream S3; wherein the stream S2 subjected to (iv) has a temperature in the range of from 150 to 400°C, more preferably in the range of from 200 to 375 °C, more preferably in the range of from 250 to 350 °C.

As to the heterogeneous hydroprocessing catalyst according to (iv), any heterogeneous hydroprocessing catalyst can be used as far as it permits to obtain a stream S3. Preferably, the heterogeneous hydroprocessing catalyst according to (iv) comprises an element of the groups 8 to 10, more preferably 9 and 10 of the periodic table of elements, more preferably an element selected from the group consisting of Ni and Co, wherein the hydroprocessing catalyst more preferably comprises Ni.

Preferably, the heterogeneous hydroprocessing catalyst according to (iv) comprises Ni in an amount, calculated as NiO, in the range of from 0.5 to 10 weight-%, more preferably in the range of from 1 to 6 weight-%, based on the weight of the hydroprocessing catalyst.

Preferably, the heterogeneous hydroprocessing catalyst according to (iv) further comprises a support for the element of the groups 8 to 10 of the periodic table of elements, wherein the support preferably is an oxidic material. Preferably, the oxidic material is one or more of alumina, silica, magnesia, zirconia, zinc oxide, calcium oxide, mixed silica-alumina, zeolite, Mo-doped alumina and titania, more preferably alumina, zeolite and silica-alumina, more preferably alumina.

Preferably, the heterogeneous hydroprocessing catalyst according to (iv) further comprises an element of the group 6 of the periodic table of elements, wherein the element of the group 6 is more preferably one or more of Mo and W. Preferably the hydroprocessing catalyst comprises in the range of from 1 to 40 weight-%, more preferably from 3 to 30 weight-%, of an oxide of said element of the group 6, more preferably Mo oxide or W oxide, based on the weight of the hydroprocessing catalyst. Preferably the heterogeneous hydroprocessing catalyst according to (iv) comprises Ni and Mo on a support, more preferably a support as defined in in the foregoing, wherein the hydroprocessing catalyst more preferably comprises Ni and Mo on alumina.

According to the present invention, the hydroprocessing catalyst, in particular the hydroprocessing catalyst comprising Ni, may comprise from 0.1 to 5 weight-%, preferably from 0.1 to 4 weight-%, more preferably from 0.1 to 3 weight phosphorus, calculated as P2Osand based on the total weight of the catalyst.

A hydroprocessing catalyst comprising Ni can be presulfidized, prior to being used, according to the following (or a similar) method: the catalyst is dried at 200 °C (temperature increase rate 1 K/min) for at least 2 h under flow of nitrogen (GHSV=2000/h) at atmospheric pressure until no water is condensed anymore downstream. Afterwards the catalyst is cooled down to 135 °C and hydrogen is fed to the reactor at a GHSV=2000/h and the reactor is pressurized (1 bar/min). After 1 h hexadecane spiked with 2 weight-% dimethyldisulfide is dosed with LHSV= ,2/h for 1 h. Afterwards temperature is increased with 0.25 K/min to 350 °C and kept for 2h. Afterwards temperature and pressure are adjusted before hexadecane solution dosing is stopped and feed is dosed.

Preferably, (iv) comprises

(iv.1 ) introducing a gas stream G2 into Z3, G2 comprising H2;

(iv.2) introducing the stream S2 obtained from (iii) into Z3;

(iv.3) bringing S2 in contact with G2 and a heterogeneous hydroprocessing catalyst comprised in Z3, obtaining a stream S3;

(iv.4) removing S3 obtained in (iv.3) from Z3.

Preferably, the gas stream G2 has a temperature in the range of 250 to 550°C, more preferably in the range of from 300 to 450 °C, more preferably in the range of from 325 to 400 °C.

Preferably the gas stream G2 is introduced at a pressure in the range of from 20 to 150 bar (abs), more preferably in the range of from 30 to 90 bar(abs), more preferably in the range of from 40 to 80 bar(abs), more preferably in the range of from 45 to 60 bar(abs).

Preferably, in Z3 the liquid hourly space velocity (LHSV) is in the range of from 0.1 to 10 m ³ /(m ³ h), more preferably in the range of from 0.15 to 5 m ³ /(m ³ h), more preferably in the range of from 0.2 to 2 m ³ /(m ³ h), wherein the LHSV is defined as the volume flow of S2 through Z3 (in m ³ /h) per volume of heterogeneous hydroprocessing catalyst comprised in Z3 (in m ³ ).

According to the present invention, G2 may preferably be recycled from the stream obtained from hydroprocessing.

Preferably, from 50 to 100 weight-%, more preferably from 70 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream G2 consists of H2. Further conceivable ranges are from 92 to 100 weight-% or from 94 to 100 weight-% or from 96 to 100 weight-% or from 98 to 100 weight-%.

According to (iv.1 ), G2 is preferably introduced continuously or semi-continuously, more preferably continuously, into Z3.

According to (iv.2), S2 is preferably introduced continuously or semi-continuously, more preferably continuously, into Z3.

Preferably, the reaction zone Z3 is comprised in a reactor, more preferably comprising n serially coupled catalyst beds B(i), i= 1 ..., n, n > 2, wherein a catalyst bed B(i) comprises a heterogeneous hydroprocessing catalyst, more preferably 2 < n < 10, more preferably 2 < n < 5; wherein B(1) is the most upstream catalyst bed and B(n) is the most downstream catalyst bed.

More preferably, (iv) comprises

(iv.1 ’) introducing a gas stream G2 into Z3, G2 comprising H2;

(iv.2’) introducing the stream S2 obtained from (iii) into Z3;

(iv.3’) n successive process stages P(i), i=1...n, wherein in P(1 )

- the gas stream G2 is introduced into a catalyst bed B(1) and brought in contact with the stream S2 obtained from (iii) and a heterogeneous hydroprocessing catalyst in B(1), obtaining a stream SP(1 ); wherein in each P(i), when i=2...n-1 , a gas stream F(i-1 ), comprising H2, is introduced into a catalyst bed B(i) and brought in contact with Sp(i-1 ) and a heterogeneous hydroprocessing catalyst in B(i), obtaining a stream Sp(i); removing Sp(i) from B(i); and wherein in P(n), a gas stream F(n-1 ) is introduced into a catalyst bed B(n) and brought in contact with Sp(n-1 ) and a heterogeneous hydroprocessing catalyst in B(n), obtaining a gas stream S3;

(iv.4’) removing S3 obtained in (iv.3’) from Z3.

Preferably, the gas stream G2 has a temperature in the range of 250 to 550°C, more preferably in the range of from 300 to 450 °C, more preferably in the range of from 325 to 400 °C.

Preferably, the gas stream G2 is introduced at a pressure in the range of from 20 to 150 bar(abs), more preferably in the range of from 30 to 90 bar(abs), more preferably in the range of from 40 to 80 bar(abs), more preferably in the range of from 45 to 60 bar(abs).

Preferably, from 50 to 100 weight-%, more preferably from 70 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream G2 consists of H2. Further conceivable ranges are from 92 to 100 weight-% or from 94 to 100 weight-% or from 96 to 100 weight-% or from 98 to 100 weight-%. Preferably, in Z3, the liquid hourly space velocity (LHSV) is in the range of from 0.1 to 10 IT ¹ , preferably in the range of from 0.1 to 5 IT ¹ , more preferably in the range of from 0.2 to 2 IT ¹ , wherein the LHSV is defined as the volume flow of S2 or Sp(i) through Z3 (in m ³ /h) per volume of heterogeneous hydroprocessing catalyst comprised in Z3 (in m ³ ).

According to (iv.1 ’), G2 is preferably introduced continuously or semi-continuously, more preferably continuously, into Z3.

According to (i v.2’), S2 is preferably introduced continuously or semi-continuously, more preferably continuously, into Z3.

Preferably, from 50 to 100 weight-%, more preferably from 70 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream F(i) consists of H2. Further conceivable ranges are from 92 to 100 weight-% or from 94 to 100 weight-% or from 96 to 100 weight-% or from 98 to 100 weight-%.

Preferably, the gas stream F(i) is introduced at a pressure in the range of from 20 to 150 bar(abs), more preferably in the range of from 30 to 90 bar(abs), more preferably in the range of from 40 to 80 bar(abs), more preferably in the range of from 45 to 60 bar(abs).

Preferably, the n serially coupled catalyst beds B(i) are fixed catalyst beds.

Alternatively, the reaction zone Z3 is comprised in a reactor comprising one catalyst bed, wherein the catalyst bed comprises a heterogeneous hydroprocessing catalyst.

In the context of the present invention, the reaction zone Z3 is preferably comprised in a continuous stirred tank reactor (CSTR) or a fixed bed reactor, more preferably in a fixed bed reactor, more preferably a trickle bed reactor.

Preferably, the stream S3 obtained from (iii) has a reduced total chlorine content of from 90 to 100 %, more preferably of from 95 to 100 %, more preferably of from 99 to 100%, more preferably of from 99.5 to 100%, compared to SO and S1 .

Preferably, the stream S3 obtained from (iv) has a total chlorine content in the range of from 0 to 50 wppm (ppm by weight), more preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 10 wppm, more preferably from 0 to 5 wppm, more preferably from 0 to 2 wppm, determined as described in Reference Example 2.1.

Preferably, the stream S3 obtained from (iv) has a chloride content in the range of from at most 40 wppm (ppm by weight), preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 1 wppm, determined as described in Reference Example 2.2. Preferably, the stream S3 obtained from (iii) has a reduced nitrogen content of from 80 to 100 %, more preferably of from 85 to 100 %, more preferably of from 90 to 100%, compared to SO, S1 and S2.

Preferably, the stream S3 obtained from (iv), more preferably after removing dissolved NH3, has a nitrogen content in the range of from 0 to 200 ppm by weight (wppm), more preferably in the range of from 0 to 100 wppm, more preferably from 0 to 50 wppm, more preferably from 0 to 10 wppm, determined as described in Reference Example 3.

Preferably, the stream S3 obtained from (iii) has a reduced sulfur content of from 80 to 100 %, more preferably of from 85 to 100 %, more preferably of from 90 to 100%, compared to SO, S1 and S2.

Preferably, the stream S3 obtained from (iv), more preferably after removing dissolved H2S, has a sulfur content in the range of from 0 to 200 ppm by weight (wppm), more preferably from 0 to 100 wppm, more preferably from 0 to 50 wppm, determined as described in Reference Example 4.

Preferably, the stream S3 obtained from (iii) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of 0 to 3 g(l2)/100 g, more preferably from 0 to 2 g(l2)/100 g, more preferably from 0 to 1 g(l2)/100 g, more preferably from 0 to 0.25 g(l2)/100 g, more preferably from 0 to 0.1 g(l2)/100 g of the stream S3 obtained from

(iv), determined as described in Reference Example 1.

Preferably, the stream S3 has a chlorine content, a nitrogen content, a sulfur content and a total amount of the one or more organic compounds comprising conjugated double bonds as defined in the foregoing.

Preferably, the stream S3 obtained from (iv), preferably prior to being subjected to (v) as defined in the following, is depleted in one or more nitrogen-containing compounds and sulfur- containing compounds compared to S2.

As mentioned in the introductory part of the invention, polymers comprised in plastic waste may contain oxygen. Such oxygen is preferably removed in the above-described hydroprocessing step.

Steps downstream of (iv)

The process of the present invention preferably further comprises, after (iii), or (iv) as defined in the foregoing,

(v) one or more of a steam cracking step, hydrocracking step, distillation, stripping, and an aqueous extraction. Preferably, the process of the present invention is a continuous or semi-continuous process, more preferably a continuous process.

Preferably, the process of the present invention consists of (i), (ii), (iii), more preferably of (i), (ii), (iii) and (iv), more preferably of (i), (ii), (iii), (iv) and (v).

The present invention further relates to a production unit for carrying out the process for purifying a pyrolysis oil according to the present invention, the unit comprising at least one reaction zone Z1 , Z1 comprising a heterogeneous hydrogenation catalyst; an inlet means for introducing SO into Z1 ; an outlet means for removing S1 from Z1 ; at least one dehalogenation zone Z2, Z2 preferably comprising a heterogeneous adsorption material or a heterogenous dehalogenation catalyst; an inlet means for introducing S1 into Z2; an outlet means for removing S2 from Z2; wherein Z1 is located upstream of Z2.

Preferably, the production unit further comprises an extraction zone to remove halides, more preferably comprising one or more halides of N-containing organic compounds, said extraction zone more preferably being arranged downstream of Z2.

Preferably, the production unit further comprises at least one reaction zone Z3, Z3 comprising a heterogeneous hydroprocessing catalyst; an inlet means for introducing S2 into Z3; an outlet means for removing S3 from Z3; wherein Z2 is located upstream of Z3.

The present invention further relates to a purified pyrolysis oil, obtainable or obtained by a process according to the present invention comprising (i), (ii) and (iii), or (i), (ii), (iii) and (iv).

When the process according to the present invention comprises (i), (ii) and (iii), the purified pyrolysis oil has preferably a chlorine content and a total amount of the one or more organic compounds comprising conjugated double bonds as defined in the foregoing for the stream S2, more preferably a chlorine content, a nitrogen content a sulfur content and a total amount of the one or more organic compounds comprising conjugated double bonds as defined in the foregoing for the stream S2.

When the process according to the present invention comprises (i), (ii), (iii) and (iv), the purified pyrolysis oil has preferably a chlorine content, a nitrogen content a sulfur content and a total amount of the one or more organic compounds comprising conjugated double bonds as defined in the foregoing for the stream S3. Preferably, the purified pyrolysis oil has a total chlorine content in the range of from 0 to 50 wppm (ppm by weight), more preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 10 wppm, more preferably from 0 to 5 wppm, more preferably from 0 to 2 wppm, determined as described in Reference Example 2.1 .

Preferably, the purified pyrolysis oil has a chloride content in the range of from at most 40 wppm (ppm by weight), preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 1 wppm, determined as described in Reference Example 2.2.

Preferably, the purified pyrolysis oil has a nitrogen content in the range of from 0 to 200 ppm by weight (wppm), more preferably in the range of from 0 to 100 wppm, more preferably from 0 to 50 wppm, more preferably from 0 to 10 wppm, determined as described in Reference Example 3.

Preferably, the purified pyrolysis oil has a sulfur content in the range of from 0 to 200 ppm by weight (wppm), more preferably from 0 to 100 wppm, more preferably from 0 to 50 wppm, determined as described in Reference Example 4.

Preferably, the purified pyrolysis oil comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of 0 to 3 g(l2)/100 g, more preferably from 0 to 2 g(l2)/100 g of the purified pyrolysis oil, determined as described in Reference Example 1.

The present invention is further illustrated by the following set of embodiments and combinations of embodiments resulting from the dependencies and back-references as indicated. In particular, it is noted that in each instance where a range of embodiments is mentioned, for example in the context of a term such as "The process of any one of embodiments 1 to 4", every embodiment in this range is meant to be explicitly disclosed for the skilled person, i.e. the wording of this term is to be understood by the skilled person as being synonymous to "The process of any one of embodiments 1 , 2, 3 and 4". Further, it is explicitly noted that the following set of embodiments represents a suitably structured part of the general description directed to preferred aspects of the present invention, and, thus, suitably supports, but does not represent the claims of the present invention.

1 . A process for purifying a pyrolysis oil, the process comprising:

(i) providing a stream SO comprising a pyrolysis oil, the pyrolysis oil comprising one or more halogenated organic compounds and one or more organic compounds comprising conjugated double bonds;

(ii) subjecting the stream SO provided in (i) to hydrogenation in at least one reaction zone Z1 containing a heterogeneous hydrogenation catalyst, obtaining a stream S1 being depleted, compared to SO, in the one or more organic compounds comprising conjugated double bonds; (iii) subjecting the stream S1 obtained from (ii) to dehalogenation in at least one dehalogenation zone Z2 downstream of Z1 , obtaining a stream S2 being depleted, compared to S1 , in the one or more halogenated organic compounds.

2. The process of embodiment 1 , wherein from 95 to 100 weight-%, preferably from 98 to 100 weight-%, more preferably from 99 to 100 weight-%, of SO consist of pyrolysis oil.

3. The process of embodiment 1 or 2, wherein the pyrolysis oil according to (i) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of from 0.1 to 75 g(l2)/100 g, preferably from 0.4 to 60 g(l2)/100 g, more preferably from 1 to 30 g(l2)/100 g of the pyrolysis oil, determined as described in Reference Example 1 .

4. The process of any one of embodiments 1 to 3, wherein the one or more organic compounds comprising conjugated double bonds comprise one or more organic compounds according to formula (I)

R ¹ R ² C ¹ =C ² R ³ -C ³ R ⁴ =X (I) wherein =X is =0, =S, =NR ⁵ , or =C ⁴ R ⁶ R ⁷ , preferably =C ⁴ R ⁶ R ⁷ .

5. The process of embodiment 4, wherein R ¹ , R ² , R ³ , R ⁴ , R ⁵ are, independently of each other, H, alkyl having from 1 to 6 carbon atoms, alkenyl having from 1 to 6 carbon atoms, or aryl having from 5 to 10 carbon atoms, more preferably H; wherein R ⁶ and R ⁷ are, independently of each other, H, alkyl having from 1 to 6 carbon atoms, alkenyl having from 1 to 6 carbon atoms, or aryl having from 5 to 10 carbon atoms, more preferably H; or wherein either R ⁴ and R ⁶ or R ⁴ and R ⁷ are linked together, thus forming, together with C ³ =C ⁴ , an aromatic ring preferably having 5 or 6 members.

6. The process of embodiment 5, wherein the one or more organic compounds according to formula (I) comprise one or more of butadiene, isoprene, dienes having 5 or 6 carbon atoms, styrene, methylstyrene, indene, substituted styrene, substituted indene, and 3- methyl-2-butenal, more preferably comprise one or more of butadiene, isoprene, dienes having 5 or 6 carbon atoms, styrene, methylstyrene, indene, and 3-methyl-2-butenal.

7. The process of embodiment 6, wherein the one or more organic compounds according to formula (I) comprise styrene.

8. The process of any one of embodiments 1 to 7, wherein the one or more halogenated organic compounds comprise one or more of mono- oligo- or polyhalogenated aromatic compounds, alkylhalides and alkenylhalides. The process of any one of embodiments 1 to 8, wherein the pyrolysis oil according to (i) has a total chlorine content in the range of from 30 to 3,000 wppm (ppm by weight), preferably from 30 to 500 wppm, more preferably from 30 to 200 wppm, determined as described in Reference Example 2.1 ; wherein the pyrolysis oil according to (i) has a chloride content of at most 40 wppm, more preferably in the range of from 0 to 30 wppm, determined as described in Reference Example 2.2. The process of any one of embodiments 1 to 9, wherein the pyrolysis oil according to (i) has a nitrogen content in the range of from 50 to 20,000 wppm (ppm by weight), preferably from 50 to 5,000 wppm, more preferably from 100 to 4,000 wppm, determined as described in Reference Example 3. The process of any one of embodiments 1 to 10, wherein the pyrolysis oil according to (i) has a sulfur content in the range of from 50 to 30,000 ppm by weight (wppm), preferably from 50 to 5,000 wppm, more preferably from 100 to 3,000 wppm, determined as described in Reference Example 4. The process of any one of embodiments 1 to 11 , wherein the pyrolysis oil is obtained from plastic waste. The process of any one of embodiments 1 to 12, wherein the stream SO is a liquid stream. The process of any one of embodiments 1 to 13, wherein the stream SO subjected to hydrogenation in (ii) has a temperature in the range of from 60 to 250 °C, preferably from 80 to 220 °C, more preferably from 100 to 200 °C. The process of any one of embodiments 1 to 14, wherein the heterogeneous hydrogenation catalyst according to (ii) comprises an element of the groups 8 to 12, preferably 8 to 10, more preferably 9 and 10, of the periodic table of elements, preferably an element selected from the group consisting of Ni, Pd and Co, more preferably from the group consisting of Ni and Pd. The process of embodiment 15, wherein the heterogeneous hydrogenation catalyst according to (ii) further comprises a support material for said element of the groups 8 to 12 of the periodic table of elements, wherein the support material is preferably selected from the group consisting of an oxidic material and carbon, wherein the oxidic material is preferably one or more of alumina, silica, magnesia, zirconia, titania, a zeolitic material, a sili- ca-alumina phosphate (SAPO) material, zinc oxide, sodium oxide, mixed silica-alumina, zeolite and calcium oxide, more preferably alumina. The process of embodiment 15 or 16, wherein the heterogeneous hydrogenation catalyst according to (ii) comprises Ni, preferably in an amount, calculated as NiO, in the range of from 0.5 to 70 weight-%, more preferably from 0.75 to 45 weight-%, more preferably from 1 to 20 weight-%, based on the total weight of the hydrogenation catalyst. The process of any one of embodiments 15 to 17, wherein the heterogeneous hydrogenation catalyst used in (ii) further comprises an element of the group 6 of the periodic table of elements, wherein the element of the group 6 is preferably one or more of Mo and W, more preferably Mo; wherein the hydrogenation catalyst preferably comprises from 1 to 40 weight-%, more preferably from 2 to 35 weight-%, more preferably from 3 to 30 weight-% of said element of the group 6, based on the total weight of the hydrogenation catalyst. The process of embodiment 17 or 18, wherein the heterogeneous hydrogenation catalyst according to (ii) comprises Ni and Mo supported on a support material, preferably a support material as defined in embodiment 16, wherein the hydrogenation catalyst preferably comprises Ni and Mo supported on alumina. The process of embodiment 15 or 16, wherein the heterogeneous hydrogenation catalyst according to (ii) comprises Pd, preferably in an amount, calculated as elemental Pd, in the range of from 0.01 to 5 weight-%, preferably from 0.1 to 1 weight-%, more preferably from 0.15 to 0.8 weight-%, based on the total weight of the catalyst. The process of embodiment 20, wherein the heterogeneous hydrogenation catalyst according to (ii) further comprises a promoter, the promoter preferably being one or more of an element of the groups 10 and 11 of the periodic table of elements, preferably one or more of Cu, Au, Ag, and Pt, more preferably one or more of Ag and Pt, more preferably Ag. The process of embodiment 21 , wherein the atomic ratio of the element of groups 8 to 12 of the periodic table, preferably Pd, relative to the promoter is in the range of from 0.1 :1 to 10:1 , preferably from 2:1 to 7:1 , more preferably from 2.5:1 to 6:1. The process of any one of embodiments 20 to 22, wherein the heterogeneous hydrogenation catalyst according to (ii) comprises Pd supported on a support material, preferably a support material as defined in embodiment 16, wherein the support material is preferably alumina or carbon, more preferably alumina. The process of any one of embodiments 1 to 23, wherein the heterogeneous hydrogenation catalyst according to (ii) is in the form of extrudates, pellets, rings, spherical particles or spheres, preferably spherical particles or extrudates. The process of any one of embodiments 1 to 24, wherein (ii) comprises

(11.1) introducing a gas stream GO into Z1 , the gas stream comprising H2;

(11.2) introducing the stream SO into Z1 ; (11.3) bringing SO in contact with GO and the heterogeneous hydrogenation catalyst comprised in Z1 , obtaining a stream S1 being depleted, compared to SO, in the one or more organic compounds comprising conjugated double bonds;

(11.4) removing S1 from Z1.

(11.5) optionally removing a gas stream G1 from Z1 , G1 comprising H2.

26. The process of embodiment 25, wherein the gas stream GO has a temperature in the range of 100 to 250 °C, preferably from 120 to 220 °C, more preferably from 140 to 200 °C.

27. The process of embodiment 25 or 26, wherein the gas stream GO is introduced at a pressure in the range of from 10 to 100 bar(abs), preferably from 15 to 90 bar(abs), more preferably from 20 to 80 bar(abs), more preferably in the range of from 20 to 55 bar(abs).

28. The process of any one of embodiments 25 to 27, wherein from 70 to 100 weight-%, preferably from 80 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream GO consists of H2.

29. The process of any one of embodiments 25 to 28, wherein according to (ii.1 ), GO is introduced continuously or semi-continuously, preferably continuously into Z1 , and wherein according to (ii.2), SO is introduced semi-continuously or continuously, preferably continuously, into Z1.

30. The process of embodiment 29, wherein GO is introduced into Z1 according to (ii.1 ) for a period At prior to introducing SO into Z1 according to (ii.2).

31 . The process of embodiment 30, wherein during At, GO is brought in contact with the heterogeneous hydrogenation catalyst comprised in Z1 , wherein GO has a temperature in the range of 50 to 250 °C, preferably from 120 to 220 °C, more preferably from 140 to 200 °C.

32. The process of any one of embodiments 25 to 31 , wherein in Z1 , the liquid hourly space velocity (LHSV) is in the range of from 0.2 to 10 m ³ /(m ³ h), preferably in the range of from 0.3 to 5 m ³ /(m ³ h), more preferably in the range of from 0.5 to 2 m ³ /(m ³ h), wherein the LHSV is defined as the volume flow of SO through Z1 (in m ³ /h) per volume of heterogeneous hydrogenation catalyst comprised in Z1 (in m ³ ).

33. The process of any one of embodiments 1 to 32, wherein the reaction zone Z1 is comprised in a continuous stirred tank reactor (CSTR) or a fixed bed reactor, preferably in a fixed bed reactor, wherein the fixed bed reactor is preferably a trickle bed reactor.

34. The process of any one of embodiments 1 to 33, wherein according to (ii), two or more reaction zones Z1 are employed which are arranged serially and/or in parallel, or wherein one single reaction zone Z1 is employed according to (ii), preferably according to (ii), one single reaction zone Z1 is employed. The process of any one of embodiments 1 to 35, preferably of any embodiment as far as being dependent on embodiment 3, wherein the stream S1 obtained from (ii) and subjected to dehalogenation in (iii) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of from 0 to 3 g(l2)/100 g, preferably from 0 to 2 g(l2)/100g, more preferably from 0 to 1 g(l2)/100 g, more preferably from 0 to 0.25 g(l2)/100 g, more preferably from 0 to 0.1 g(l2)/100 g of the stream S1 , determined as described in Reference Example 1. The process of any one of embodiments 1 to 35, wherein the stream S1 subjected to dehalogenation in (iii) has a temperature in the range of from 150 to 450°C, preferably from 200 to 400°C, more preferably from 250 to 350 °C. The process of any one of embodiments 1 to 36, wherein the dehalogenation zone Z2 according to (iii) comprises, preferably is an adsorption zone, preferably comprising a heterogeneous adsorbent material suitable for adsorbing halide comprised in at least one of the one or more halogenated organic compounds, preferably in all of the one or more halogenated organic compounds. The process of embodiment 37, wherein the heterogeneous adsorbent material according to (iii) comprises one or more of a carbon-containing adsorbent material and an alumi- num-containing adsorbent material, preferably an aluminum-containing adsorbent material; wherein the carbon-containing adsorbent material is preferably a carbon-containing molecular sieve, more preferably activated charcoal; wherein the aluminum-containing adsorbent material is preferably an alumina, an aluminum-containing molecular sieve, a silicoaluminophosphate, a silica-alumina hydrate or a hydrotalcite; wherein the aluminum-containing molecular sieve is preferably an alumina, an aluminosilicate, preferably having a molar ratio of Si:AI, calculated as SiC^AhOs, in the range of from 2:1 to 10:1 , more preferably from 2:1 to 4:1 ; wherein the silica-alumina hydrate preferably has weight ratio AhO3:SiO2 in the range of from 1 :1 to 10:1 , more preferably from 1 :1 to 2:1 ; wherein the hydrotalcite is preferably an aluminum and magnesium containing hydrotalcite, more preferably an aluminum-magnesium hydroxycarbonate, preferably having a MgO:AhO3 weight ratio in the range of from 63:37 to 70:30; wherein the heterogeneous adsorbent material more preferably comprises the hydrotalcite; wherein the heterogeneous adsorbent material according to (iii) preferably comprises an element of the groups 1 , 2, 11 and 12. 39. The process of any one of embodiments 1 to 38, wherein the heterogeneous adsorbent material according to (iii) comprises particles characterized by a particle size distribution having a D50 value in the range of from 1 to 6,500 micrometers, preferably from 2 to 2,000 micrometers, more preferably from 8 to 500 micrometers, more preferably from 10 to 50 micrometers or from 3 to 9 micrometers, the D50 particle size being determined as described in Reference Example 5.

40. The process of any one of embodiments 1 to 39, wherein the heterogeneous adsorbent material according to (iii) has an average pore volume in the range of from 0.1 to 5 ml/g, preferably in the range of from 0.15 to 2 ml/g, the average pore volume being determined as described in Reference Example 6.

41 . The process of any one of embodiments 1 to 40, wherein the heterogeneous adsorbent material according to (iii) has a BET specific surface area in the range of from 50 to 1 ,000 m ² /g, preferably in the range of from 100 to 900 m ² /g, more preferably in the range of from 150 to 600 m ² /g, the BET specific surface area being determined as described in reference Example 7.

42. The process of any one of embodiments 1 to 41 , wherein (iii) comprises

(111.1 ) introducing a gas stream G1 into Z2 preferably being an adsorption zone, preferably a gas stream comprising one or more of hydrogen and nitrogen, more preferably hydrogen;

(111.2) introducing the stream S1 obtained from (ii) into Z2;

(111.3) bringing S1 in contact with G1 and a heterogeneous adsorbent material comprised in Z2, obtaining a stream S2 being depleted, compared to S1 , in the one or more halogenated organic compounds;

(111.4) removing S2 from Z2.

43. The process of embodiment 42, wherein the gas stream G1 has a temperature in the range of 250 to 500°C, preferably in the range of from 300 to 400 °C.

44. The process of embodiment 42 or 43, wherein the gas stream G1 is introduced at a pressure in the range of from 1 to 100 bar(abs), preferably in the range of from 5 to 80 bar(abs), more preferably in the range of from 10 to 50 bar(abs).

45. The process of any one of embodiments 42 to 44, wherein in Z2, the liquid hourly space velocity (LHSV) is in the range of from 0.2 to 10 IT ¹ , preferably in the range of from 0.3 to 5 IT ¹ , more preferably in the range of from 0.5 to 2 IT ¹ .

46. The process of any one of embodiments 42 to 45, wherein from 90 to 100 weight-%, preferably from 95 to 100 weight-%, more preferably from 98 to 100 weight-%, of the gas stream G1 consists of H2; or wherein from 90 to 100 weight-%, preferably from 95 to 100 weight-%, more preferably from 98 to 100 weight-%, of the gas stream G1 consists of nitrogen.

47. The process of any one of embodiments 42 to 46, wherein according to (iii.1 ), G1 is introduced continuously or semi-continuously, preferably continuously, into Z2 and wherein according to (iii.2) S1 is introduced continuously or semi-continuously, preferably continuously, into Z2; wherein the adsorption zone Z2 is preferably comprised in a continuous stirred tank reactor (CSTR) or a fixed bed reactor, preferably in a fixed bed reactor, more preferably a trickle bed reactor, the reactor preferably comprising an adsorption bed comprising the heterogeneous adsorbent material.

48. The process of any one of embodiments 1 to 47, wherein, according to (iii), two or more reaction zones Z2 are employed which are arranged serially and/or in parallel, wherein preferably, one single reaction zone Z2 is employed according to (iii).

49. The process of any one of embodiments 1 to 48, wherein the stream S2 obtained from (iii) has a total chlorine content in the range of from 0 to 200 wppm (ppm by weight), preferably from 0 to 160 wppm, more preferably from 0 to 130 wppm, more preferably from 0 to 120 wppm, determined as described in Reference Example 2.1 ; wherein the stream S2 obtained from (iii) has a chloride content of at most 40 wppm (ppm by weight), preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 1 wppm, determined as described in Reference Example 2.2.

50. The process of any one of embodiments 1 to 49, preferably of any embodiment as far as being dependent on embodiment 3, wherein the stream S2 obtained from (iii) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of 0 to 3 g(l2)/100 g, preferably from 0 to 2 g(l2)/100 g, more preferably from 0 to 1 g(l2)/100 g, more preferably from 0 to 0.25 g(l2)/100 g, more preferably from 0 to 0.1 g(l2)/100 g of the stream S2, determined as described in Reference Example 1.

51 . The process of any one of embodiments 1 to 50, wherein the stream S2 obtained from (iii) has a nitrogen content in the range of from 50 to 20,000 ppm by weight wppm, preferably from 50 to 5,000 wppm, more preferably from 100 to 4,000 wppm, determined as described in Reference Example 3; and wherein the stream S2 obtained from (iii) has a sulfur content in the range of from 50 to 30,000 ppm by weight (wppm), preferably from 50 to 5,000 wppm, more preferably from 100 to 3,000 wppm, determined as described in Reference Example 4.

52. The process of any one of embodiments 1 to 36, wherein the dehalogenation zone Z2 according to (iii) comprises, preferably is a catalytic zone, preferably comprising a hetero- geneous dehalogenation catalyst, said catalyst comprising one or more catalytically active elements of groups 8 to 12 of the periodic system of elements.

53. The process of any one embodiments 1 to 52, wherein the stream S2 obtained from (iii), prior to being subjected to hydroprocessing according to (iv) as defined in embodiment 54, is subjected to extraction, preferably using an aqueous extraction medium, obtaining a stream S2 being depleted in one or more dissolved halides comprised in S2 obtained from the dehalogenation zone Z2, said halides preferably comprising one or more halides of N- containing organic compounds.

54. The process of any one of embodiments 1 to 53, further comprising

(iv) subjecting the stream S2 obtained from (iii) to hydroprocessing in at least one reaction zone Z3 downstream of Z2, Z3 comprising a heterogeneous hydroprocessing catalyst; obtaining a stream S3; wherein the stream S2 subjected to (iv) has a temperature in the range of from 150 to 400°C, preferably in the range of from 200 to 375 °C, more preferably in the range of from 250 to 350 °C.

55. The process of embodiment 54, wherein the heterogeneous hydroprocessing catalyst used in (iv) comprises an element of the groups 8 to 10, preferably 9 and 10 of the periodic table of elements, preferably an element selected from the group consisting of Ni and Co, wherein the hydroprocessing catalyst more preferably comprises Ni; wherein the heterogeneous hydroprocessing catalyst according to (iv) more preferably comprises Ni in an amount, calculated as NiO, in the range of from 0.5 to 10 weight-%, more preferably in the range of from 1 to 6 weight-%, based on the weight of the hydroprocessing catalyst.

56. The process of embodiment 55, wherein the heterogeneous hydroprocessing catalyst according to (iv) further comprises a support for the element of the groups 8 to 10 of the periodic table of elements, wherein the support preferably is an oxidic material; wherein the oxidic material preferably is one or more of alumina, silica, magnesia, zirconia, zinc oxide, calcium oxide, mixed silica-alumina, zeolite, Mo-doped alumina and titania, more preferably alumina, zeolite and silica-alumina, more preferably alumina.

57. The process of embodiment 55 or 56, wherein the heterogeneous hydroprocessing catalyst according to (iv) further comprises an element of the group 6 of the periodic table of elements, wherein the element of the group 6 is preferably one or more of Mo and W; wherein the hydroprocessing catalyst preferably comprises in the range of from 1 to 40 weight-%, more preferably from 3 to 30 weight-%, of an oxide of said element of the group 6, preferably Mo oxide or W oxide, based on the weight of the hydroprocessing catalyst.

58. The process of embodiment 57, wherein the heterogeneous hydroprocessing catalyst according to (iv) comprises Ni and Mo on a support, preferably a support as defined in embodiment 53, wherein the hydroprocessing catalyst preferably comprises Ni and Mo on alumina.

59. The process of any one of embodiments 53 to 58, wherein (iv) comprises (iv.1 ) introducing a gas stream G2 into Z3, G2 comprising H2;

(iv.2) introducing the stream S2 obtained from (iii) into Z3;

(iv.3) bringing S2 in contact with G2 and a heterogeneous hydroprocessing catalyst comprised in Z3, obtaining a stream S3;

(iv.4) removing S3 obtained in (iv.3) from Z3.

60. The process of embodiment 59, wherein the gas stream G2 has a temperature in the range of 250 to 550°C, preferably in the range of from 300 to 450 °C, more preferably in the range of from 325 to 400 °C.

61 . The process of embodiment 59 or 60, wherein the gas stream G2 is introduced at a pressure in the range of from 20 to 150 bar (abs), preferably in the range of from 30 to 90 bar(abs), more preferably in the range of from 40 to 80 bar(abs), more preferably in the range of from 45 to 60 bar(abs).

62. The process of any one of embodiments 59 to 61 , wherein in Z3 the liquid hourly space velocity (LHSV) is in the range of from 0.1 to 10 IT ¹ , preferably in the range of from 0.1 to 5 IT ¹ , more preferably in the range of from 0.2 to 2 IT ¹ .

63. The process of any one of embodiments 59 to 62, wherein from 50 to 100 weight-%, preferably from 70 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream G2 consists of H2.

64. The process of any one of embodiments 59 to 63, wherein according to (iv.1 ) G2 is introduced continuously or semi-continuously, preferably continuously, into Z3.

65. The process of any one of embodiments 59 to 64, wherein according to (iv.2) S2 is introduced continuously or semi-continuously, preferably continuously, into Z3.

66. The process of any one of embodiments 1 to 65, wherein the reaction zone Z3 is comprised in a reactor, preferably comprising n serially coupled catalyst beds B(i), i= 1 ..., n, n > 2, wherein a catalyst bed B(i) comprises a heterogeneous hydroprocessing catalyst, preferably 2 < n < 10, more preferably 2 < n < 5; wherein B(1 ) is the most upstream catalyst bed and B(n) is the most downstream catalyst bed.

67. The process of embodiment 66, wherein (iv) comprises

(iv.T) introducing a gas stream G2 into Z3, G2 comprising H2;

(iv.2’) introducing the stream S2 obtained from (iii) into Z3;

(iv.3’) n successive process stages P(i), i=1...n, wherein in P(1 ) - the gas stream G2 is introduced into a catalyst bed B(1) and brought in contact with the stream S2 obtained from (iii) and a heterogeneous hydroprocessing catalyst in B(1), obtaining a stream SP(1 ); wherein in each P(i), when i=2...n-1 , a gas stream F(i-1 ), comprising H2, is introduced into a catalyst bed B(i) and brought in contact with Sp(i-1 ) and a heterogeneous hydroprocessing catalyst in B(i), obtaining a stream Sp(i); removing Sp(i) from B(i); and wherein in P(n), a gas stream F(n-1) is introduced into a catalyst bed B(n) and brought in contact with Sp(n-1) and a heterogeneous hydroprocessing catalyst in B(n), obtaining a gas stream S3;

(iv.4’) removing S3 obtained in (iv.3’) from Z3.

68. The process of embodiment 67, wherein the gas stream G2 has a temperature in the range of 250 to 550°C, preferably in the range of from 300 to 450 °C, more preferably in the range of from 325 to 400 °C.

69. The process of embodiment 67 or 68, wherein the gas stream G2 is introduced at a pressure in the range of from 20 to 150 bar (abs) , preferably in the range of from 30 to 90 bar(abs), more preferably in the range of from 40 to 80 bar (abs), more preferably in the range of from 45 to 60 bar(abs).

70. The process of any one of embodiments 67 to 69, wherein from 50 to 100 weight-%, preferably from 70 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream G2 consists of H2.

71 . The process of any one of embodiments 67 to 70, wherein in Z3 the liquid hourly space velocity (LHSV) is in the range of from 0.1 to 10 IT ¹ , preferably in the range of from 0.1 to 5 IT ¹ , more preferably in the range of from 0.2 to 2 IT ¹ .

72. The process of any one of embodiments 67 to 71 , wherein according to (iv.T) G2 is introduced continuously or semi-continuously, preferably continuously, into Z3.

73. The process of any one of embodiments 67 to 72, wherein according to (iv.2’) S2 is introduced continuously or semi-continuously, preferably continuously, into Z3.

74. The process of any one of embodiments 67 to 73, wherein from 50 to 100 weight-%, preferably from 70 to 100 weight-%, more preferably from 90 to 100 weight-%, of the gas stream F(i) consists of H2.

75. The process of embodiment 74, wherein the gas stream F(i) is introduced at a pressure in the range of from 20 to 150 bar(abs), preferably in the range of from 30 to 90 bar(abs), more preferably in the range of from 40 to 80 bar (abs), more preferably in the range of from 45 to 60 bar(abs).

76. The process of any one of embodiments 67 to 75, wherein the n serially coupled catalyst bed B(i) are fixed catalyst beds.

77. The process of any one of embodiments 1 to 76, wherein the reaction zone Z3 is comprised in a continuous stirred tank reactor (CSTR) or a fixed bed reactor, preferably in a fixed bed reactor, more preferably a trickle bed reactor.

78. The process of any one of embodiments 53 to 77, wherein the stream S3 obtained from (iv) has a total chlorine content in the range of from 0 to 50 wppm (ppm by weight), preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 10 wppm, more preferably from 0 to 5 wppm, more preferably from 0 to 2 wppm, determined as described in Reference Example 2.1 ; wherein the stream S3 obtained from (iv) has a chloride content in the range of from at most 40 wppm (ppm by weight), preferably from 0 to 30 wppm, more preferably from 0 to 20 wppm, more preferably from 0 to 1 wppm, determined as described in Reference Example 2.2.

79. The process of any one of embodiments 53 to 78, wherein the stream S3 obtained from (iv), preferably after removing dissolved NH3, has a nitrogen content in the range of from 0 to 200 ppm by weight (wppm), preferably in the range of from 0 to 100 wppm, more preferably from 0 to 50 wppm, more preferably from 0 to 10 wppm, determined as described in Reference Example 3.

80. The process of any one of embodiments 53 to 79, wherein the stream S3 obtained from

(iv), preferably after removing dissolved H2S, has a sulfur content in the range of from 0 to 200 ppm by weight (wppm), preferably from 0 to 100 wppm, more preferably from 0 to 50 wppm, determined as described in Reference Example 4.

81 . The process of any one of embodiments 1 to 80, preferably of any embodiment as far as being dependent on embodiment 3, wherein the stream S3 obtained from (iii) comprises the one or more organic compounds comprising conjugated double bonds in a total amount in the range of 0 to 3 g(l2)/100 g, preferably from 0 to 2 g(l2)/100 g, more preferably from 0 to 1 g(l2)/100 g, more preferably from 0 to 0.25 g(l2)/100 g, more preferably from 0 to 0.1 g(l2)/100 g of the stream S3 obtained from (iv), determined as described in Reference Example 1.

82. The process of any one of embodiments 1 to 81 , further comprising, after (iii), or (iv) as defined in any one of embodiments 53 to 81 ,

(v) one or more of a steam cracking step, hydrocracking step, distillation, stripping, and an aqueous extraction. 83. The process of any one of embodiments 1 to 82, being a continuous or semi-continuous process, preferably a continuous process.

84. The process of any one of embodiments 1 to 83, consisting of (i), (ii), (iii), preferably (i), (ii), (iii) and (iv), more preferably (i), (ii), (iii), (iv) and (v).

85. A production unit for carrying out the process for purifying a pyrolysis oil according to any one of embodiments 1 to 84, the unit comprising at least one reaction zone Z1 , Z1 comprising a heterogeneous hydrogenation catalyst; an inlet means for introducing SO into Z1 ; an outlet means for removing S1 from Z1 ; at least one dehalogenation zone zone Z2, Z2 preferably comprising a heterogeneous adsorption material or a heterogenous dehalogenation catalyst; an inlet means for introducing S1 into Z2; an outlet means for removing S2 from Z2; wherein Z1 is located upstream of Z2; and preferably an extraction zone to remove halides, preferably comprising one or more halides of N-containing organic compounds, said extraction zone preferably being arranged downstream of Z2.

86. The production unit of embodiment 85, further comprising at least one reaction zone Z3, Z3 comprising a heterogeneous hydroprocessing catalyst; an inlet means for introducing S2 into Z3; an outlet means for removing S3 from Z3; wherein Z2 is located upstream of Z3.

87. A purified pyrolysis oil, obtainable or obtained by a process according to any one of embodiments 1 to 53 or according to any one of embodiments 54 to 81 .

Examples

Reference Example 1 Measurement of the total amount of the one or more organic compounds comprising conjugated double bonds

The diene content is determined by UOP326-17. In this procedure dienes are reacted with maleic anhydride (MA) and the consumption of MA is determined (by titration of the remainder MA). It can be expressed as g(l2)/100g(sample) or alternatively as g(MA)/100g(sample). The unit can be interconverted by multiplying the MA-value (MAV) by a factor 2.59 to obtain the value expressed with g(l2)/100g(sample) corresponding to the molar weight of I2 and MA. Accordingly, 1wt% Styrene or 0.52 wt.-% Butadiene correspond to 0.94g(MA)/100g or 2.43g(l2)/100g. Reference Example 2 .1 Measurement of total chlorine content (wppm)

The sample is filtered with a 0.45pm syringe filter before analysis. The chlorine content is determined by combustion of the respective sample at 1050°C. Resulting combustion gases, i.e., hydrogen chloride, are led into a cell in which coulometric titration is performed.

Reference Example 2.2 Measurement of chloride content (wppm)

The sample is filtered with a 0.45pm syringe filter before analysis. The chloride content is determined by ion chromatography. Apparatus: Ion chromatograph 850 Professional (Metrohm) (Pre column: Metrosep A Supp4/5 S-Guard and Analytical column: Metrosep A Supp 5 250/4; Flow: 0.7 mL/min; Column temperature: 30°C; Detector temperature: 40°C; Inject volume: 25 pL; Suppressor MSM HC Rotor A). As Eluant: 3.2 mmol/L Na2CO3 ; 1.0 mmol/L NaHCO3 and as Suppressor regenerant: 50 mmol/L sulfuric acid were used.

Sample preparation: 0.2 g - 0.4 g of the sample were weighed and dissolved in 10 mL toluene. For analyte extraction, 10 mL deionized water were added. After centrifugation, the aqueous phase was extracted and analyzed. Samples with a concentration below the limit value of the method were spiked with 20 pg/L chloride standard solution (corresponding to a limit value of 1 mg/kg chloride in the sample) to check the recovery rate.

Reference Example 3 Measurement of N content (wppm)

The nitrogen content is determined by combustion of the respective sample at 1000°C. NO contained in resulting combustion gases reacts with ozone so that NO2* is formed. Relaxation of excited nitrogen species is detected by chemiluminescence detectors according to ASTM D4629 (N). Calibration range is from 0.5 wppm to 50 wppm. Samples with higher concentrations are diluted with xylene to be in calibration range.

Reference Example 4 Measurement of S content (wppm)

The sulfur content is determined by combustion of the respective sample at 1000°C.

Sulfur dioxide which is contained in resulting combustion gases is excited by UV (ultraviolet) light. Light which is emitted during relaxation is detected by UV fluorescence detectors according to ASTM D5453 (S). Calibration range is from 0.5 wppm to 50 wppm. Samples with higher concentrations are diluted with xylene to be in calibration range.

Reference Example 5 Particle size (D50)

The D50 particle size was determined by optical methods or by an air sieve, for example by various instruments, namely, Cilas Granulometer 1064 supplied by Quantachrome, Malvern Mastersizer or Luftstrahlsieb (air sieve) supplied by Alpine. Reference Example 6 Determination of the average pore volume

Pore volume can be derived from BET measurements (for micro and mesopores) or alternatively Hg porosimetry (for macropores). The Determination of Pore Volume and Area Distributions in Porous Substances. I. Computations from Nitrogen Isotherms JACS 1951 (73) 373-380 E.P. Barret, L.G. Joyner, P.P. Halenda.

Reference Example 7 Determination of the BET specific surface area

The BET surface area of the adsorbent material is measured by using an instrument supplied by Quantachrome (Nova series) or by Micromeritics (Gemini series). The method entails low temperature adsorption of nitrogen at the BET region of the adsorption isotherm.

Reference Example 8 Determination of the styrene content

GC method with a nonpolar, 100% dimethylpolysiloxane phase column and FID-detector. Final column temperature and inlet temperature are 330°C and 320°C, respectively. Integrated area signal of Styrene as ratio of all integrated peaks times 100% is Area%. Area% roughly correlates with wt.%.

Reference Example 9 Determination of the total acid number (TAN)

The total acid number was determined by titration with KOH according to ASTM D3242.

Reference Example 10 Reactor loading and test setup

All reaction steps of the examples were conducted in reactors with inner diameter of 10mm and were operated in trickle bed mode (downflow). The reactor (80cm in length) is loaded with corundum (WSK F46; commercial corundum) from the bottom such that the lower 25cm are filled with inert corundum (cooling zone). On top of this the 30cm long catalyst or adsorbents bed is placed from 25 to 55cm and in this zone the reaction temperature is maintained. In case of shaped catalysts as P-doped NiMo-catalyst and E-157 SDU catalyst, the void spaces of the catalyst bed are filled with corundum (WSK F46; commercial corundum) as well. On top of the catalyst resp. adsorbents bed corundum (WSK F46; commercial corundum) is filled from 55 to 80cm. In this corundum zone the feed is preheated to the reaction temperature whereas in the lower corundum filled zone the product stream is cooled from the reactor temperature down to the trace heating temperature.

Example 1 Process for purifying a pyrolysis oil according to the present invention

A feed stream SO comprising a pyrolysis oil having a MAV of 9.01 g(l2)/100g , a styrene content of 6.1 Area% determined by GC, a total chlorine content of 560 wppm, a chloride content of 1 wppm, a nitrogen content of about 3260 wppm, a sulfur content of about 2630 wppm, a total acid number (TAN) of 3.69 mg(KOH/g(feed) and a density of 0.8653 g/ml was subjected to hydrogenation in a reactor comprising an activated Pd-catalyst (Catalyst E-157 SDU 1/8” commercially available from BASF: 0.7wt.-%Pd on 1/8” alumina beads) with H2 at 50 bar and at a temperature of 195 °C. The Pd-catalyst was activated in hydrogen flow with GHSV=1000/h at ambient pressure with heating to 195°C at a ramp of 0.5K/min. After 12h at 195°C GHSV was reduced to 500/h and pressure was increased to 50 bar within 1 h and kept for 2h before addition of SO started with LHSV=1/h. The resultant feed stream S1 had a MAV reduced by about 80% to 1.75 g(l2)/100g and the styrene content was reduced to 1 .1 Area% (> 80% conversion), the N- and S-contents were not changed compared to SO. Also, the chlorine content was basically unaltered (cf. Table 1).

Table 1 Chlorine and chloride contents for SO and S1

The stream S1 was then subjected to dechlorination in a reactor comprising a Cl-adsorbent (hydrotalcite: aluminum-magnesium hydroxycarbonate powder having a MgO:Al2O3 weight ratio of 70:30) in the presence of H2 at 350 °C and at a pressure of 50 bar. Before use, the adsorbent was compacted, then crushed and sieved to an average particle size of 500-1000 micrometers. Further, it was calcined at 450°C for 5h in air and equilibrated in ambient air overnight. Prior to entering S1 in the reactor, the obtained Cl-adsorbent was dried at 100°C and 200°C at a gas hourly space velocity (GHSV) =2000/h in nitrogen under ambient pressure for 1 hour each while ramping temperature with 1 K/min. At 200°C, gas was switched from N2 to H2 and the pressure was increased to 50 bar within 1 hour. After the pressure was attained, the GHSV was reduced to 475/h and the reactor was heated to 350°C with 1 K/min. Once 350 °C were attained, S1 was introduced in the reactor at a liquid hourly space velocity (LHSV) =0.95/h. The resultant product stream was analyzed as shown in Table 2.1 and had a reduced total chlorine content (at the beginning of the operation reduced by about 90 %).

Table 2.1 Chlorine and chloride contents for S1 and product stream from dechlorination at various time on stream (TOS) The intermediate products (S2) within the course of the reaction were combined to form the feed S2.

The stream S2 was further washed with an equivalent volume of water and analyzed (see Table 2.2). The N-, S-contents were not changed compared to S1.

Table 2.2 Chlorine and chloride contents of combined products (S2) before and after washing

Further, the (washed) stream S2 was subjected to hydroprocessing in a reactor comprising an activated P-doped NiMo-catalyst supported on alumina (4.75wt%NiO, 19.57wt% MoOs, 2.89wt% P2O5 and 72.79% support resp. AI2O3; prepared according to US 4,409,131) in the presence of H2 at 50 bar and temperatures ranging from 350-395 °C. Before the P-doped NiMo- catalyst was activated by being dried at ambient pressure and GHSV=2000/h with N2 upon heating with 1 K/min and dwelled for 2h at 200°C. Thereafter, for the catalyst activation, temperature was reduced to 135°C and the atmosphere was switched to H2. After 1 h, the pressure was increased to 50 bar within 1 hour. After another hour the dosing of sulfidation feed was started with LHSV=2/h and the GHSV was adjusted to 1000/h. Sulfidation feed consists of hydrocarbon fluids (commercial mixture of boiling point range of 176-209°C, such as Varsol™ 60) spiked with 2 wt.% DM DS (Dimethyldisulfide). After 2 h, the temperature was further increased with 0.25K/min up to 350°C. Then, the temperature was kept constant for 1 h, LHSV was reduced to 1/h and GHSV was reduced to 500/h before the addition of S2 started in the reactor at a LHSV=1/h.

After 73h TOS, the temperature was increased to 395°C and the system was operated another 49h resulting compared to comparative example 1 in longer operation times at both temperatures without any indication of increasing pressure drop. Afterwards LHSV and GHSV were reduced by 50% for another day before pressure was increased to 110 bar. LHSV-reduction and pressure increase did increase the N- and S-conversion further.

The resultant feed stream S3 had in average a total chlorine content of below 2 wppm, a chloride content of below 1 wppm (see Table 2.3) and also greatly reduced N- and S-contents of from 3264 wppm up to about 31 wppm for N content and of from 2630 wppm up to about 47 wppm for S content (see Table 3).

Table 2.3 Chlorine and chloride contents before and after hydroprocessing

Table 3 Nitrogen and Sulfur contents before and after hydroprocessing

Comparative Example 1 Process for purifying a pyrolysis oil not according to the present invention (representative of FR 3 107 530)

A feed stream SO comprising a pyrolysis oil as the one in Example 1 was subjected to hydrogenation in a reactor comprising an activated Pd-catalyst (Catalyst E-157 SDU 1/8” commercially available from BASF: 0.7 wt.-%Pd on 1/8” alumina beads) with H2 at 50 bar and at a tempera- ture of 195 °C. The Pd-catalyst was activated in hydrogen flow with GHSV=1000/h at ambient pressure with heating to 195°C at a ramp of 0.5K/min. After 12h at 195°C GHSV was reduced to 500/h and pressure was increased to 50 bar within 1 h and kept for 2h before addition of SO started with LHSV=1/h. The resultant feed stream S1 had a MAV reduced by about 80% to 1.75 g(l2)/100g and the styrene content was reduced to 1.1 Area% (> 80% conversion), the N-, S- contents were not changed compared to SO, i.e. the N-content was of about 3260 wppm and the S-content was of about 2630 wppm. The chlorine and chloride contents are in Table 4.

Table 4 Chlorine and chloride contents for SO and S1

Further, the stream S1 was directly subjected to hydroprocessing in a reactor comprising an activated P-doped NiMo-catalyst supported on alumina (4.75wt%NiO, 19.57wt% MoOs, 2.89wt% P2O5 and 72.79% support resp. AI2O3; the catalyst was prepared according to a process as described in US 4,409,131) in the presence of H2 at 50 bar and temperatures ranging from 350-395 °C. The catalyst activation was made as in Example 1 . The temperature in the reactor was kept constant after the catalyst activation for 1 h, LHSV was reduced to 0.95/h and GHSV was reduced to 475/h before the addition of S1 started in the reactor with LHSV=1/h.

After 48h TOS, the reactor temperature was increased to 395°C but after about 70h TOS the pressure drop in the reactor system increased above 10 bar so that the feed dosing was stopped and the reactor was cooled down. A first intermediate pressure drop increase was observed from 55 to 60h TOS.

To demonstrate that such pressure increase was due to plugged downstream tubes, the tube was exchanged as explained in the following. Indeed, to enable access to the downstream section, the system was depressurized and purged with nitrogen. A 1 m long thin tubing heated to 40°C (as the complete following downstream section) with 1 ,5mm ID was located after the transfer line from the reactor (180°C, 4mm ID) and before the gas/liquid separator. This tube was exchanged. This change of tube permitted to decrease the pressure drop across the reactor system substantially so that it was not noticeable anymore.

Furthermore, 5mg of a solid were removed from the tube. By XRD, NH4CI was detected in the solid as the main crystalline phase as can be seen in Figure 1.

After this tube exchange, the reactor was restarted with sulfidation feed and hydrogen and heated with 1 K/min to 350°C. At this temperature the original feed was dosed again allowing another 15 h of hydroprocessing operation before again a pressure drop increase was observed so that the feed dosing was stopped after about 100h TOS. Despite intermediate system plugging, the catalyst still revealed a relatively good performance after restart with N-conversion of >90% (cf. Table 5). For the combined product streams a total chlorine content <2 wppm and a chloride content <1 wppm was determined suggesting that converted Cl is not present in the liquid product stream in dissolved form and noticeable amounts or is present but as particulate matter which is filtered off upon sample preparation for Cl-analysis.

Table 5 Nitrogen and Sulfur contents before and after hydroprocessing

Table 6 Comparison pressure drop after hydroprocessing of prehydrogenated and dechlorinated feed (Ex. 1) and after hydroprocessing of prehydrogenated feed (C. Ex. 1)

The pressure drop of the reactor is externally determined with 100mln(N2)/min at ambient outlet pressure. Resulting values are given in Table 6.

Compared to the process of the present invention, the aforementioned process (representative of FR 3 107 530) which does not comprise a dechlorination step between the hydrogenation and hydroprocessing steps presents high pressure drop increase rates during the reaction course which do not allow stable operation unless water is injected downstream which causes serious corrosion issues. In addition, also increased pressure drop of the reactors after the reaction (210 mbar vs. 71 mbar with the inventive process; cf. Table 6) was observed. Therefore, the process of the present invention presents a great improvement compared to this known process in view of more stable operation regarding increasing pressure drop and thus allowing increased TOS, which permits to obtain an improved oil production in terms of stability and duration.

Comparative Example 2 Process for purifying a pyrolysis oil according to the present invention

As opposed to Example 1 , the process was started with a dechlorination step. A feed stream SO comprising a pyrolysis oil as the one in Example 1 was subjected to dechlorination in a reactor comprising a Cl-adsorbent (hydrotalcite: aluminum-magnesium hydroxycarbonate powder having a MgdAhOs weight ratio of 70:30) in the presence of H2 at 350 °C and at a pressure of 50 bar. Before use, the adsorbent was compacted, then crushed and sieved to an average particle size of 500-1000 micrometers. Further, it was calcined at 450°C for 5h in air and equilibrated in ambient air overnight. Prior to entering SO in the reactor, the obtained Cl-adsorbent was dried at 100°C and 200°C at a gas hourly space velocity (GHSV) =2000/h in nitrogen under ambient pressure for 1 hour each while ramping temperature with 1 K/min. At 200°C, gas was switched from N2 to H2 and the pressure was increased to 50 bar within 1 hour. After the pressure was attained, the GHSV was reduced to 475/h and the reactor was heated to 350°C with 1 K/min. Once 350 °C were attained, SO was introduced in the reactor at a liquid hourly space velocity (LHSV) =0.95/h.

After 165 h TOS, the pressure drop in the reactor system increased above 10 bar so that the feed dosing had to be stopped and the reactor was cooled down, purged with toluene and dried with nitrogen. The pressure drop of the reactor is externally determined with 100mln(N2)/min at ambient outlet pressure resulting in 1102 mbar whereas the reactor's pressure drop before testing was determined with 37 mbar. The chlorine and chloride contents in the obtained stream S2 are in Table 7.

Table 7 Chlorine and chloride contents before and after dechlorination

Table 8 Comparison pressure drop after dechlorination alone or after pre-hydrogenation + dechlorination

After the reaction the reactors were unloaded. The different parts of the reactor loading were separated into the inert corundum material on top (preheating zone) and below (cooling zone) the adsorbents zone and the adsorbent. Especially within the preheating zone color gradients were observed across the length. Therefore, samples were homogenized before analysis. Samples were heated in air up to 700°C and the resultant mass loss was recorded by thermo- gravimetric analysis in conjunction with differential scanning calorimetry (device STA 449 F3 Jupiter from company Netzsch). For all samples obtained from comparative example 2 the exothermic mass loss was increased which is indicative of increased coke formation. The relative increase of coke formation to example 1 is most pronounced in the preheating zone with roughly 100% increase of the mass loss.

In Figures 2a and 2b the coloring of the corundum from the preheating zone can be compared showing a darker color for dechlorination only (C. Ex. 2 - 2a). After heating in air, samples did decolorize to white as in the fresh state.

Table 9 Comparison of the mass loss in wt.% upon heating corundum of the preheating zone in air after dechlorination or after pre-hydrogenation + dechlorination

As both dechlorination experiments in example 1 and comparative example 2 were operated for the same duration, conditions and with the same adsorbents and inert corundum and showed large differences in pressure drop of the reactors afterwards and mass loss in heating up the material of the preheating zone in air it can be derived that the coke build-up in the dechlorination step is accelerated if feed is not prehydrogenated.

Therefore, based on this example, it has been demonstrated that a hydrogenation step is mandatory prior to the dechlorination step. Indeed, starting with dechlorination directly results in extremely fast increase of pressure loss by coking which is not acceptable for industrial plant operation. In addition, the dechlorination efficiency (cf. Table 2.1 and 7) is increased if the feed is prehydrogenated before dechlorination.

Brief description of the figures

Figure 1 is the analysis by XRD of the solid obtained in Comp. Ex. 1 contained in the tube . Reflexes of crystalline NH4CI are marked with asterisks.

Figure 2 a. is a picture of the corundum of the preheating zone for Comparative Example 2 (black powder). b. is a picture of the corundum of the preheating zone for Example 1 (powder lighter than in a.).

Cited literature

- WO 2017/083018 A1

- FR 3 103 822