Self-quenching gasification process

Abstract

A process for simultaneously producing a mixture of hydrocarbons from a carbonaceous feedstock and recovering heat, the process comprising the steps of a) providing, as a fluidization steam, a stream of steam via at least one gas inlet to a bottom section of a fluidized bed reactor; b) contacting said stream of steam with a solid bed material thereby transferring heat from said stream of steam to said bed material such that a fluidized stream comprising steam and a heated bed material are obtained; c) feeding a carbonaceous feedstock into said fluidized bed reactor; d) contacting said carbonaceous feedstock with said fluidized stream comprising steam and a heated fluidized bed material in a reaction zone of said fluidized bed reactor thereby forming a fluidized bed, e) allowing a fluid stream comprising steam and a mixture of hydrocarbons in gaseous form to leave the fluidized bed at the top; (f) allowing said fluid stream of steam and hydrocarbons in gaseous form to reside in a free gas volume above the fluidized bed in order to allow for the decay of radicals present in said stream of hydrocarbons in gaseous form; (g) withdrawing a gaseous product mixture comprising a mixture of gaseous hydrocarbons from said fluidized bed reactor and transferring at least part thereof directly to a heat exchanger in order to recover heat; wherein in step c) said carbonaceous feedstock is fed to the reactor via at least one inlet on the top of the reactor; and wherein the fluidized stream has a fluidization velocity in a range of from 2 to 10 times the minimum fluidization velocity as calculated using the Ergun equation.

Description

SELF-QUENCHING GASIFICATION PROCESSField of the inventionThe present invention relates to a self-quenching cracking process for converting a carbonaceous feedstock, in particular plastic and / or organic waste, to valuable hydrocarbon compounds. BackgroundThe production of valuable hydrocarbon compounds, in particular monomers such as ethylene and propylene for polymer synthesis, via chemical recycling of waste is gaining more and more importance in view of global developments.Commercial scale processes for the gasification of carbonaceous feedstock are currently under development. For example, WO2022038316A1 discloses a gasification process that serves to convert plastic and organic waste to valuable hydrocarbon compounds. The process involves oxygen free gasification at elevated temperatures of 600 to 900 °C. Furthermore, the prior art implies that there are yield increases of certain product fractions, propylene in particular, when the product stream is cooled, or quenched, rapidly.However, rapid quenching requires advanced process equipment installed in immediate connection to the reactor. Designs are often vulnerable to material wear due to heat stresses resulting from the process set-up but also repeated patterns of start-up and shut down. This challenge is well known to petro-chemical operators, e.g. steam crackers. Moreover, to achieve the heat transfer coefficients required for rapid cooling, a phase change is typically needed, meaning the quench coolers incorporate high pressure steam generation systems. This type of equipment constitutes significant capital expenditure due to material requirements and introduces a serious process safety risk due to possible material failure.Furthermore, given the experience of conventional steam cracking, it can reasonably be assumed that the equipment, unless skilfully designed and operated, will be prone to fouling. This is due to the fact that unsaturated hydrocarbon compounds tend to aggregate and form an insulating layer, or even plugging, thus decreasing efficiency of the cooler and by extension causing lower propylene yields throughout the run. The issue can also reasonably be foreseen to increase a waste recycling environment where contaminant substances e.g. metals can catalyse formation fouling substances.Object of the inventionIt is an object of the invention to overcome the above drawbacks and, in particular, to provide a gasification process providing high yields of valuable hydrocarbon compounds that does not require external quenching. Summary of the inventionIt has now been found that the above objective can be achieved by a process for simultaneously producing a mixture of hydrocarbons from a carbonaceous feedstock and recovering heat, the process comprising the steps of a) providing,   as a fluidization steam, a stream of steam via at least one gas inlet to a bottom section of a fluidized bed reactor; b) contacting said stream of steam with a solid bed material thereby transferring heat from said stream of steam to said bed material such that a fluidized stream comprising steam and a heated bed material are obtained; c) feeding a carbonaceous feedstock into said fluidized bed reactor; d) contacting said carbonaceous feedstock with said fluidized stream comprising steam and a heated fluidized bed material in a reaction zone of said fluidized bed reactor thereby forming a fluidized bed, e) allowing a fluid stream comprising steam and a mixture of hydrocarbons in gaseous form to leave the fluidized bed at the top; (f) allowing said fluid stream of steam and hydrocarbons in gaseous form to reside in a free gas volume above the fluidized bed in order to allow for the decay of radicals present in said stream of hydrocarbons in gaseous form; (g) withdrawing a gaseous product mixture comprising a mixture of gaseous hydrocarbons from said fluidized bed reactor and transferring at least part thereof directly to a heat exchanger in order to recover heat; wherein in step c) said carbonaceous feedstock is fed to the reactor via at least one inlet on the top of the reactor; and wherein the fluidized stream has a fluidization velocity in a range of from 2 to 10 times the minimum fluidization velocity as calculated using the Ergun equationwherein: ρp is the particle density in kg / m³; ρf is the density of the fluid;g is gravity; µf is the dynamic viscosity of the fluid; Vmf is the minimum fluidization velocity; εmf is the void fraction (porosity) of the bed; Φs is sphericity of the bed material particles; andDp is the average particle size of the bed material in mm.A key feature of the invention is to remove the need for rapid quenching without causing any detrimental effects on the product distribution. Instead of connecting the reactor to a rapid quench cooler, the gasification vessel is designed to quench the reactions automatically already in the reactor. The cooling of the product stream thus becomes less time critical allowing for a simpler design.The reactor contains a steam fluidised stream of solid bed material into which the feed is introduced. Upon forming, product gases flow out of the fluidised bed to a particle free volume above the bed. The gasifier is specifically designed to let the radical pool in the gas phase product decay which removes the driver of the reactions by for example minimised heat transfer into the gaseous volume. The specific heat of incoming steam is small compared to the reaction heat needed to drive the endothermic reactions that consume desirable fractions.The invention removes the need for quench coolers in immediate connection to the reactor, while maintaining the yield of light product gases (C3 and lighter) at the cost of aromatics formation. This is achieved by creating a gas volume above the fluidised bed in the reactor. The gas volume is designed as to ensure minimal contact between the gas and heat from any solid surface. Without wanting to be bound by theory, it is assumed that the radical pool in the product stream decays rapidly and stops the reactions, thus preserving light product fractions of interest such as e.g. propylene. Heat recovery from the product stream can thus focus on heat integration rather than an immediate need to quench the reaction.Short description of the FiguresFigure 1: Schematics of the experimental set-up.  Detailed description of the inventionThe present invention relates to a process for simultaneously producing a mixture of hydrocarbons from a carbonaceous feedstock and recovering heat, the process comprising the steps of a) providing, as a fluidization steam, a stream of steam via at least one gas inlet to a bottom section of a fluidized bed reactor; b) contacting said stream of steam with a solid bed material thereby transferring heat from said stream of steam to said bed material such that a fluidized stream comprising steam and a heated bed material are obtained; c) feeding a carbonaceous feedstock into said fluidized bed reactor; d) contacting said carbonaceous feedstock with said fluidized stream comprising steam and a heated fluidized bed material in a reaction zone of said fluidized bed reactor thereby forming a fluidized bed, e) allowing a fluid stream comprising steam and a mixture of hydrocarbons in gaseous form to leave the fluidized bed at the top; (f) allowing said fluid stream of steam and hydrocarbons in gaseous form to reside in a free gas volume above the fluidized bed in order to allow for the decay of radicals present in said stream of hydrocarbons in gaseous form; (g) withdrawing a gaseous product mixture comprising a mixture of gaseous hydrocarbons from said fluidized bed reactor and transferring at least part thereof directly to a heat exchanger in order to recover heat; wherein in step c) said carbonaceous feedstock is fed to the reactor via at least one inlet on the top of the reactor; andwherein the fluidized stream has a fluidization velocity in a range of from 2 to 10 times the minimum fluidization velocity as calculated using the Ergun equation    wherein: ρp is the particle density in kg / m³; ρf is the density of the fluid;g is gravity; µf is the dynamic viscosity of the fluid; Vmf is the minimum fluidization velocity; εmf is the void fraction (porosity) of the bed; Φs is sphericity of the bed material particles; andDp is the average particle size of the bed material in mm.Preferably, in the process of the present invention, the gas residence time – which corresponds to the time period for which said fluid stream of steam and hydrocarbons in gaseous form is allowed to reside in a free gas volume above the fluidized bed (step f) – is greater than 1 second, preferably from 1 to 8 seconds, more preferably from 2 to 7 seconds and even more preferably from 4.2 to 6.2 seconds.It is to be noted that in conventional steam cracking processes, the gas residence time window is much narrower—typically around 0.2 seconds—before quenching is required. In fact extending the gas residence time in a conventional steam cracking process would lead to significant issues such as coking and fouling in the cracker tubes and the Transfer Line Exchanger (TLE). However, the present invention demonstrates that this time can be safely extended in the fluidized bed without encountering noticeable coking or fouling problemsPreferably, the process of the present invention is carried out in a dual fluidized bed reactor (DFB) comprising a fluidized bed gasifier, at least one gas analyzer being fluidly connected to a product gas outlet of said fluidized bed gasifier and at least one control unit for adjusting the composition of feedstock; wherein said fluidized bed gasifier comprises at least one, preferably at least two inlets for introducing said carbonaceous feedstock. Preferably, the cross section of the freeboard is equal to the cross section of the fluidized bed. Preferably, the height of the freeboard is at least equal to the height of the fluidized bed. Preferably, feeding is done via on-bed feeding, that is feedstock is added to the fluidized bed gasifier at a position above the fluidized bed. Preferably, at least one inlet for introducing said carbonaceous feedstock to the fluidized bed reactor assembly comprises an extruder. A schematic drawing of an exemplary DFB system is depicted in Figure 1, wherein said DFB system comprises (1) a combustor (or regenerator); (2) fuel feed for the combustor; (3) a cyclone; (4) a particle distributor; (5) a first loop seal to the gasifier; (6) a gasifier; (7) a second loop seal to the combustor; (8) fuel feed for the gasifier. Moreover, a sampling point (X) as well as return points of the second loop seal (crossed circle) are included in Figure 1.Moreover, details regarding a reactor set-up for conventional steam cracking can be found, for example, in Ullmann’s Encyclopedia of Industrial Chemistry (DOI: 10.1002 / 14356007).Preferably, in the process of the present invention, no external quenching takes place. Preferably, the temperature in the reaction zone is in a range of from 500 to 900°C, preferably in a range of from 600 to 850°C. Preferably, the process is carried out at sub-atmospheric pressure, preferably at -0.5 to -5 kPa gauge pressure, more preferably at -1 to -2 kPa gauge pressure. Preferably, the bed material is transferred to and regenerated in a second fluidized bed reactor which is fluidly connected with said first fluidized bed reactor. The type of carbonaceous feedstock is not particularly limited and may be selected, for example, from mixed (solid) waste, mixed plastic waste, polyolefins, a plastic waste derived pyrolysis oil of aliphatic character with a high degree of unsaturated substances, vegetable oil, animal fat, a reject fraction of mechanical recycling, a reject fraction of paper recycling, (pre-) sorted plastic waste, oils and waxes of synthetic or biogenic origin and biomass.Preferably, at least part of said carbonaceous feedstock is in a liquid state at ambient temperature and pressure.Preferably, said carbonaceous feedstock comprises one or more fluid pyrolysis products, more preferably one or more fluid pyrolysis products obtained from plastic waste that has not been subject to hydrotreatment operations.More preferably, the carbonaceous feedstock comprises at least one of mixed plastic waste, biomass and liquid aliphatic and / or naphthenic hydrocarbons.The bed material to be used in the process described herein is not particularly limited. However, preferably, the bed material comprises an inert bed material.Preferably, the fluidization steam consists of steam only; viz. no additional hydrogen and / or oxygen gas is introduced. Ideally, the heat of reaction is provided exclusively by the fluidized bed. The feedstock is converted upon contact with the hot bed material. As soon as the converted feedstock leaves the surface of the bed material, the convective heat transfer is no longer available to sustain the cracking reactions. It is believed that if these parameters are met, the gas residence time in the reactor becomes irrelevant.Preferably, the flow rate of the fluidization steam is adjusted such that the fluidized bed is operated in bubbling regime. Typically, for bubbling regime, the fluidization velocity is 2 to 10 times the minimum fluidization velocity. In other words, the flow rate of the fluidization steam is preferably in a range of from 100 to 300 kg / h, more preferably in a range of from 120 to 280 kg / h and even more preferably in a range of from 150 to 260 kg / h. The minimum fluidization velocity may be calculated using the Ergun equation: wherein: ρp is the particle density in kg / m³; ρf is the density of the fluid;g is gravity; µf is the dynamic viscosity of the fluid; Vmf is the minimum fluidization velocity; εmf is the void fraction (porosity) of the bed; Φs is sphericity of the bed material particlesDp is the average particle size of the bed material in mm.Likewise, the terminal velocity can be calculated (Vmax) can be calculated as follows:  Preferably, the fluidization velocity is in a range of from 2 to 8 times the minimum fluidization velocity, more preferably in a range of from 2 to 7 and even more preferably in a range of from 3 to 5 times the minimum fluidization velocity. Preferably, a first temperature of the product gas, T1, is measured 5 cm above the fluidized bed in the fluidized bed reactor and a second temperature of the product gas, T2, is measured 80 cm above the fluidized bed in the fluidized bed reactor and said second temperature T2 is at least 40°C, more preferably at least 50°C and even more preferably at least 55°C lower than said first temperature T1. Not wanting to be bound by theory, it is pointed out that in the reactor system, no external heating source is available in the freeboard, which is a region of a free gas volume above the fluidized bed inside the reactor chamber. The expressions “freeboard” and “free gas volume above the fluidized bed” may be used interchangeably herein. Therefore, the temperature gradient observed along the reactor's height is primarily due to heat losses through the reactor walls. This cooling effect is believed to contribute to the decay of free radicals in the reactor.Additionally, the rapid decrease in the partial pressure of hydrocarbons as they leave the surface of the bed material is assumed to play a significant role in free radical decay. In the freeboard, the estimated partial pressure of hydrocarbons is significantly lower than in the fluidized bed itself. Preferably, the partial pressure of hydrocarbons in the freeboard is in a range of from 0.15 to 0.5, more preferably of from 0.20 to 0.40 bar / bar (e.g. partial pressure relative to a total pressure of the reactor of 1 bar). However, within the fluidized bed itself, particularly near the bed material, the partial pressure is much higher than in the freeboard since the hydrocarbons produced within the fluidized bed do not mix uniformly with the fluidizing steam until they exit the fluidized bed. The bubbling fluidization regime, characterized by distinct bubbles of gas rising through the bed, leads to an uneven distribution of hydrocarbons and steam, resulting in regions of high hydrocarbon concentration near the bed material.Further, the cross section of the freeboard is, preferably, equal to the cross section of the fluidized bed. The height of the freeboard may be adjusted depending on how much of a temperature loss of the product gas mixture is required before the product gas exits the reactor. Preferably, the height of the freeboard is at least equal to the height of the bed material in the reactor.Consequently, it is believed that the decay of free radicals can be attributed to a combination of factors: the lack of a heat source in the area of the freeboard and the reduction in partial pressure of hydrocarbons as they move away from the high partial pressure zones within the bed.Preferably, said fluidized bed reactor is operated in bubbling regime.ExamplesDFB Industrial scale experimentsThe industrial scale experiments were performed in the Chalmers DFB system, which consists of a 12 MWth circulated fluidized bed (CFB) combustor coupled to a 2-4 MWth bubbling fluidized bed (BFB) gasifier. This configuration allows to obtain the heat needed in the gasification side by the recirculation of the bed material that comes from the combustor. To perform the steam gasification experiments, a flow of 150 kg / h of steam was used in the gasifier as fluidization media. In this case, the feeding was done in a continuous way through an extruder that has to functions: (1) to melt the feedstock allowing a more steady and homogeneous feeding to the gasifier; and (2) act as sealing preventing air to enter in the gasifier. The feeding is done via on-bed feeding. For a better view of the system, the schematic of the set-up is provided in Figure 1.To quantify the total dry gas produced during the experiments a small flow of helium was added in the gasifier as a tracer gas (35 lN / min), similar to what is done at lab scale. In this case, a raw gas stream is continuously sampled, that is used for both the permanent gases and the condensable hydrocarbons (tars). To analyze the raw gas composition, a slipstream of the dragged raw gas sampled was passed through a hot ceramic filter, cooled down and scrubbed in isopropanol to remove the condensable hydrocarbons. This cold and dry stream was then analyzed in a micro-GC (Varian CP-4900). This micro-GC has two channels and uses Poraplot Q and MS5Å columns, with He and Ar as carrier gases, respectively. The micro-GC takes a point-injection (10-30 ms injection time) of the dry and tar-free raw gas every 3 minutes, generating a new chromatogram from each injection. The micro-GC is calibrated every week with five concentration levels that cover the range of the expected concentrations. The species analyzed are: H2, He, CO, CO2, CH4, C2H2, C2H4, C2H6, C3H6, C3H8 and N2. The results of the gas composition are the average of the chromatograms taken over a period of stable operation (i.e., when the gasification temperature and the fuel flow were stable). During this stable measurement the temperature in the gasifier varied in a range of ±3 °C. To measure the tar species, the solid-phase adsorption method was used in the same way that was presented in the previous section. In this case, a set of 4 amines was taken during the stable operation. After eluation, the resulting liquid was analyzed in a BRUKER 430 GC-FID. Each sample was analyzed 3 times, and the results presented are the average of the values obtained in the three-repeat analysis for the 4 different samples. The steam cracker operating conditions used in the examples are given in Table 1; details regarding reactor dimensions and bed material are given in Tables 2 and 3. Molten HDPE is used as the feedstock and introduced through an extruder fed opening. The particle stream moves through the reactor, providing heat to the steam and reactions which produce a range of hydrocarbon compounds from the HDPE. The reactor thus has the function of a high temperature pyrolyser. The hydrocarbons and the steam leave the fluidised bed in their gaseous form at the top (thus, entering the freeboard) whereas the sand particles pass through another liquid seal and are transferred back to the regenerator. The cross section of the freeboard is equal to the cross section of the fluidized bed.The influence of the gas residence time in the gasifier was investigated by varying the flow rate of fluidisation steam between 150 and 250 kg / h. Gasification temperature was maintained at a constant level. The resultant product distribution of permanent gases exhibited very small variation as a result; see Table 4. Specifically, table 4 provides the product yield in mol / kgf of the various product species. It can thus be concluded that rapid external quenching of the gasified product stream is not necessary to safeguard product yields for this design.Table 1Operating conditionsIE1IE2IE3Steam flow to the cracker (kg / h)151201251Steam flow to the cracker (m3 / s)0.2000.2660.332Superficial gas velocity of steam, u (m / s)0.660.891.11Gasifier temperature, TG (°C)810810808Product gas temperature 5 cm above the fluidized bed, T1 (°C)761759758Product gas temperature 80 cm above the fluidized bed, T2 (°C)703702700Feeding rate polyethylene (kg / h)91.291.291.2Steam dynamic viscosity (g / cm-s)0.00038510.00038510.0003851Steam density (kg / m3)0.210.210.21Product gas flow (m3 / s)0.080.080.08Gas residence time (s)6.195.024.22Partial pressure of hydrocarbons in freeboard (bar / bar)*0.40.30.24* bar / bar: partial pressure relative to a total pressure of the reactor of 1 bar Table 2Reactor Dimensions Reactor cross sectional area (m2)0.3Reactor volume above the fluidized bed (m3)1.75 Table 3Bed Material Silica SandDp (average particle size, mm)0.2Particle density ρp (kg / m3)2660 Minimum fluidization velocity for all three operating conditions (calculated via Ergun equation) at IE1, IE2 and IE3: Vmf = 0.016 m / sThe superficial gas velocities achieved in IE1, IE2 and IE3 are greater than 0.016 m / s (refer to the table above). Thus, it is evident that the steam cracker is in fluidization mode for each of the inventive examples. The terminal velocity for IE1, IE2 and IE3 (calculated as described herein above): ut = 1.52 m / sThe superficial gas velocities achieved in IE1, IE2 and IE3 are lower than 1.52 m / s (refer to the table above). Thus, the bed material is not entrained out of the steam cracker. The fluidization regime achieved during IE1, IE2 and IE3 can be determined from the general flow regime diagram for the whole range of gas-solid contacting from percolating packed beds to lean pneumatic transport of solids; adapted from Grace (J.R. Grace, Can. J. Chem. Eng., 64, 353 (1986).) Table 4  IE1IE2IE3HDPE flowkgf / h91.291.291.2Steam flowkg / h150200250H2mol / kgf5.145.505.42CH4mol / kgf6.84 6.726.59COmol / kgf0.320.340.32CO2mol / kgf1.701.831.81C2H4mol / kgf11.0110.8310.96C2H2mol / kgf0.130.120.12C2H6mol / kgf0.890.870.89C3H6mol / kgf3.433.293.50C3H8mol / kgf0.140.130.14BTSX*mol / kgf0.921.061.01*total content of benzene, toluene, styrene and xylenes. Overall, it is evident from the above data that bed material particles above the fluidized bed are present in a very low amount which prevents convective heat transfer to the reactor volume above the gas phase. Moreover, it can be seen from the temperatures TG, T1 and T2 that the product gas leaving the bed material loses heat as it leaves the reactor (TG > T1 > T2) which results in self-quenching of the cracking reactions.In other words, the cracker effluent can go without quenching for an extended period of time after exiting the fluidized bed as compared to conventional steam cracking processes.

Claims

1. A process for simultaneously producing a mixture of hydrocarbons from a carbonaceous feedstock and recovering heat, the process comprising the steps of:a) providing, as a fluidization steam, a stream of steam via at least one gas inlet to a bottom section of a fluidized bed reactor; b) contacting said stream of steam with a solid bed material thereby transferring heat from said stream of steam to said bed material such that a fluidized stream comprising steam and a heated bed material are obtained; c) feeding a carbonaceous feedstock into said fluidized bed reactor; d) contacting said carbonaceous feedstock with said fluidized stream comprising steam and a heated fluidized bed material in a reaction zone of said fluidized bed reactor thereby forming a fluidized bed, e) allowing a fluid stream comprising steam and a mixture of hydrocarbons in gaseous form to leave the fluidized bed at the top; (f) allowing said fluid stream of steam and hydrocarbons in gaseous form to reside in a free gas volume above the fluidized bed in order to allow for the decay of radicals present in said stream of hydrocarbons in gaseous form; (g) withdrawing a gaseous product mixture comprising a mixture of gaseous hydrocarbons from said fluidized bed reactor and transferring at least part thereof directly to a heat exchanger in order to recover heat; wherein in step c) said carbonaceous feedstock is fed to the reactor via at least one inlet on the top of the reactor; andwherein the fluidized stream has a fluidization velocity in a range of from 2 to 10 times the minimum fluidization velocity as calculated using the Ergun equation wherein:ρp is the particle density in kg / m³; ρf is the density of the fluid;g is gravity; µf is the dynamic viscosity of the fluid; Vmf is the minimum fluidization velocity; εmf is the void fraction (porosity) of the bed; Φs is sphericity of the bed material particles; andDp is the average particle size of the bed material in mm.

2. The process according to claim 1, wherein no external quenching takes place.

3. The process according to claim 1, wherein the temperature in the reaction zone is in a range of from 500 to 900°C.

4. The process according to claim 3, wherein the temperature in the reaction zone is in a range of from 600 to 850°C.

5. The process according to claim 1, wherein process is carried out at sub-atmospheric pressure.  6. The process according to claim 5, wherein the process is carried out at -0.5 to -5 kPa gauge pressure.

7. The process according to claim 6, wherein the process is carried out at -1 to -2 kPa gauge pressure.

8. The process according to claim 1, wherein the bed material is transferred to and regenerated in a second fluidized bed reactor which is fluidly connected with said first fluidized bed reactor.

9. The process according to claim 1, wherein the carbonaceous feedstock comprises at least one of mixed plastic waste, biomass and liquid aliphatic and / or naphthenic hydrocarbons. 10. The process according to claim 1, wherein the fluidization velocity s in a range of from 2 to 8 times the minimum fluidization velocity.  11. The process according to claim 10, wherein the fluidization velocity s in a range of from 2 to 7 times the minimum fluidization velocity.  12. The process according to claim 11, wherein the fluidization velocity s in a range of from 3 to 5 times the minimum fluidization velocity.  13. The process according to claim 1, wherein the bed material comprises an inert bed material.  14. The process according to claim 1, wherein the flow rate of the fluidization steam is in a range of from 100 to 300 kg / h.  15. The process according to claim 14, wherein the flow rate of the fluidization steam is in a range of from 120 to 280 kg / h.  16. The process according to claim 15, wherein the flow rate of the fluidization steam is in a range of from 150 to 260 kg / h.  17. The process according to claim 1, wherein a first temperature of the gaseous product mixture, T1, is measured 5 cm above the fluidized bed in the fluidized bed reactor and a second temperature of the gaseous product mixture, T2, is measured 80 cm above the fluidized bed in the fluidized bed reactor and said second temperature T2 is at least 40°C lower than said first temperature T1.  18. The process according to claim 17, wherein said second temperature T2 is at least 50°C lower than said first temperature T1.  19. The process according to claim 18, wherein said second temperature T2 is at least 55°C lower than said first temperature T1.  20. The process according to claim 1, wherein said fluidized bed reactor is operated in bubbling regime. 21. The process according to claim 1, wherein in step f) said fluid stream of steam and hydrocarbons in gaseous form is allowed to reside in a free gas volume above the fluidized bed for more than 1 second. 22. The process according to claim 21, wherein in step f) said fluid stream of steam and hydrocarbons in gaseous form is allowed to reside in a free gas volume above the fluidized bed for 1 to 8 seconds. 23. The process according to claim 22, wherein in step f) said fluid stream of steam and hydrocarbons in gaseous form is allowed to reside in a free gas volume above the fluidized bed for 2 to 7 seconds. 24. The process according to claim 23, wherein in step f) said fluid stream of steam and hydrocarbons in gaseous form is allowed to reside in a free gas volume above the fluidized bed for 4.2 to 6.2 seconds. 

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