EA040945B1 - EXTRACTION SYSTEM FOR HIGH PRESSURE PROCESSING SYSTEM - Google Patents

Description

The field of technology to which the invention belongs

The present invention relates to the field of separation systems for use in high pressure continuous processing systems, in particular to recovery systems for recovering liquid organic compounds and/or homogeneous catalysts from the separated aqueous phase product of high pressure continuous processing systems for the conversion of carbonaceous materials. such as biomass.

State of the art invention

Numerous applications of continuous high pressure processes exist or are under development or in the early stages of commercialization. Examples of such processes are hydrothermal and solvothermal processes, for example for the production of hydrocarbons such as transport fuels, lubricants or specialty chemicals and gases from carbonaceous materials such as biomass.

High pressure conversion process products typically contain a pressurized mixture of liquid hydrocarbon compounds; a gas phase containing carbon dioxide, carbon monoxide, hydrogen, C ₁ -C ₄ hydrocarbons; an aqueous phase containing liquid organic compounds of the aqueous phase and dissolved salts, and optionally suspended solids such as inorganic substances and/or coal and/or unconverted carbonaceous material, depending on the particular carbonaceous material being processed and the particular processing conditions.

In the field of oil production, various separation methods are known. In the field of application of such hydrocarbons obtained from carbonaceous material using hydrothermal or solvothermal methods, information on the separation is limited. Hydrocarbons obtained in this way will have some characteristics similar to fossil hydrocarbons, but will subsequently differ in other areas. The hydrocarbons produced in this way, compared to fossil oils, will generally be more polarized, will have a high viscosity due to the relatively high oxygen content, and will often have a density close to that of water. The use of conventional separation methods known from fossil oil applications on the hydrocarbons thus obtained has shown that the hydrocarbons after such separation contain too much water and/or too much inorganic matter for many applications.

Typically, the product stream from the high pressure reforming process is depressurized to ambient and cooled to a temperature below the boiling point of water to allow subsequent separation into separate phases. However, although various methods have been generally proposed for separating individual phases, including solvent extraction (Downie (WO 2014/197928)), distillation (Downie (WO 2014/197928)), cyclones, e.g. hydrocyclones (Iversen (US 921317), Humfreys (WO 2008AU00429), Annee, (EP 0204354), Van de Beld (EP 1184443),), filtration (Iversen (WO 2015/092773), Iversen (US 921317 B2), Annee (EP 0204354), Downie (WO 2014 /197928), Iversen (WO 2006/117002)), decantation (Yokoyama (US 4935567), Modar (WO 81/00855)), centrifugation (Iversen (WO 2015/092773), Iversen (US 921317), Iversen (WO 2006 /117002), Annee (EP 0204354)), membrane separation (Modar (WO 81/00855), Iversen (WO 2006/117002)), only limited details have been described in the prior art regarding equipment design, separation conditions and operation .

For continuous processing, water must be removed from the process in the same amount as it is added to the process with carbonaceous material(s), catalysts, etc. The aqueous phase resulting from such separation processes typically also contains liquid organic compounds of the aqueous phase as well as dissolved salts such as homogeneous catalysts in the form of potassium and/or sodium salts and/or suspended solids, as well as other components, and requires treatment to meet environmental wastewater standards. In addition to environmental concerns, the liquid organics of the aqueous phase represent a loss of carbon that reduces the liquid hydrocarbon yield of the oil phase. Elliot et al. (US 9755828) employs a combined hydrothermal liquefaction and catalytic hydrothermal gasification system to increase total carbon yield, in which aqueous phase liquid organic compounds are reduced by hydrothermal gasification and converted into a medium BTU gas product that can be used for heating in the process. Further purification is suggested by recycling the aqueous phase and/or the solids fraction to a growth stage such as algae production. However, although the method of Elliott et al. and increases the overall yield of carbon, this is achieved through the by-product, and the yield of the target liquid hydrocarbon product of the oil phase remains unchanged. In addition, in the method of Elliott et al. nothing is said about the recovery of homogeneous catalysts in the form of potassium and sodium.

It is desirable to recover both liquid organics of the aqueous phase and homogeneous catalysts such as potassium and sodium from the aqueous phase for efficiency and economic reasons, and there is very little information in the prior art on suitable systems for such recovery and recycling to the preparation step. raw materials.

Iversen (US Patent Application No. 15/787,393) describes a recovery process in which aqueous phase liquid organic compounds and/or homogeneous catalysts are recovered from the aqueous phase using evaporation and/or distillation techniques.

While this provides some degree of recovery of some desirable components, there are other components that may require purification, in particular the aqueous liquid organic phase.

A common problem with such prior art separation systems is that the separated oil product often contains too high levels of water and inorganic matter, which limits the quality of the oil (hydrocarbons) and its further use, for example, in catalytic upgrading processes in transportation fuels, lubricants or specialized chemicals.

A common problem with such prior art separation systems is that the aqueous phase often contains too high levels of accumulated contaminants such as, for example, chlorides, which can adversely affect the process and process equipment, and as such, can directly or indirectly affect on the yield obtained in the process, the quality of the resulting product and/or the service life of the process equipment.

Accordingly, improved and more efficient separation schemes for purifying/reducing contaminants such as chlorides from the aqueous phase are desired.

Purpose of the invention

Thus, it is an object of the present invention to provide an improved separation and purification system, as well as a method of operating such a system, which at least partially recovers liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and/or sodium before they are reintroduced into the stage preparation of raw suspension.

Description of the invention

According to one aspect of the present invention, the object of the invention is achieved by a process for separating and purifying products from a high pressure processing system for processing a raw mixture containing carbonaceous material(s) at a pressure of 150 bar (15 MPa) to 400 bar (40 MPa) and temperature from 300 to 430°C in the presence of homogeneous catalysts in the form of potassium and/or sodium in a concentration of at least 0.5 wt.% and liquid organic compounds in a concentration of 5 to 40 wt.% for a predetermined time, with the formation as a result of the converted feed mixture, where the converted feed mixture is cooled to a temperature in the range of 50 to 250°C and its pressure is reduced to a pressure in the range of 1-150 bar (0.115 MPa), and where the converted feed mixture is separated into a gas phase containing dioxide carbon, hydrogen and methane, an oil phase containing liquid organic compounds of the oil phase, and an aqueous phase containing liquid organic compounds. aqueous phase organic compounds, dissolved salts, where liquid organic compounds of the aqueous phase and dissolved homogeneous catalysts in the form of potassium and/or sodium are at least partially removed from said aqueous phase, resulting in a first aqueous phase stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, and a second aqueous phase stream depleted in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, and where the first aqueous phase stream is at least partially recycled to said feed mixture to obtain at least parts of said liquid organic compounds and homogeneous catalysts in the feed mixture, and where, additionally, before recycling said first recycle stream to the feed mixture, from said aqueous phase enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, they give an outlet stream, wherein the outlet stream is additionally processed at one or more ion exchange stages and said one or more ion exchange stages contain one or more ion exchange resins contained in one or more fixed beds in a parallel arrangement with shut-off valves before and after each layer, whereby at least one ion exchange layer is connected and one or more ion exchange layers can be transferred at least offline for cleaning, thereby allowing continuous operation.

By diverting such an outlet stream from the first aqueous phase stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, it is avoided that impurity elements such as chloride accumulate in the aqueous phase as a result of said recirculation. While other impurity elements, such as polyvalent metal ions, are less soluble in the aqueous phase and can be removed from the process as solids, this is not the case for chlorides, which have high water solubility and further enhance corrosion.

In yet another preferred embodiment, the liquid organics in the feed mixture also contain recycled liquid organics in the oil phase.

As a rule, the mass ratio of said outlet stream withdrawn from the first aqueous phase stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium to the total aqueous phase stream supplied to said recovery system is in the range 0.01-0.5, for example in the range of 0.02-0.4, preferably the mass ratio of said discharged outlet stream to the total aqueous phase flow is in the range of 0.03-0.25, for example in the range of 0.04-0 ,15.

By diverting the outlet stream in the above weight ranges, it is possible to control the concentration of chlorides in the aqueous phase to acceptable values.

In a preferred embodiment, the amount of effluent withdrawn is chosen such that a chloride concentration in the feed mixture is less than 600 ppm by weight, eg less than 400 ppm by weight; preferably less than 200 ppm by weight, for example less than 100 ppm by weight.

In a preferred embodiment of the present invention, the effluent is further treated with one or more ion exchange steps.

According to a preferred embodiment, the one or more ion exchange stages comprise one or more ion exchange resins contained in one or more fixed beds in a parallel arrangement with check valves before and after each bed, whereby at least one ion exchange bed is connected, and at least one ion exchange layer is offline.

Advantageously, the ion exchange resins in said ion exchange step comprise a chloride selective resin.

Advantageously, the chloride concentration in the first aqueous phase is less than 250 ppm by weight, for example less than 200 ppm by weight; preferably, the concentration of chloride in the first aqueous phase is less than 150 ppm by weight, for example less than 100 ppm by weight.

The outlet stream, in accordance with a preferred embodiment of the invention, may be filtered to remove suspended particles prior to entering said ion exchange step(s).

In accordance with yet another preferred embodiment of the present invention, the ion exchange bed(s) is further provided with a valve mechanism to regenerate/clean said ion exchange resins by backflowing and/or backwashing the cleaning fluid while offline.

In a preferred embodiment, the cleaning fluid contains demineralized water.

The pH value at the inlet of the ion exchange stage(s) according to a preferred embodiment of the present invention is maintained in the range of 8 to 14, for example in the range of 9 to 14, preferably the pH at the inlet of the ion exchanger is in the range of 10 to 13.5.

Inlet pH maintenance may, in accordance with an embodiment of the present invention, be accomplished by measuring the pH of the outlet stream before entering the ion exchange step(s), and ultimately adding a base, such as sodium hydroxide, to the outlet stream prior to entering the ion exchange step(s). ) ion exchange or addition prior to the effluent treatment step, for example by adding a base such as sodium hydroxide in the recovery step.

The aqueous phase entering the recovery system of the present invention typically contains aqueous phase liquid organic compounds having a boiling point less than water and aqueous phase liquid organic compounds having a boiling point greater than water.

The aqueous phase liquid organic compounds recovered and introduced into the raw material preparation step of the present invention typically contain one or more components selected from one or more of the following groups:

a) ketones such as acetone, propanones, butanones, pentanones, pentenones, cyclopentanones such as 2,5-dimethylcyclopentanone, cyclopentenones, hexanones and cyclohexanones, for example 3methylhexanone, quinones;

b) alcohols and polyalcohols such as methanol, ethanol, propanols, butanols, pentanols, hexanols, heptanols, octanols such as 2-butyl-1-octanol, hydroquinones, benzene diols;

c) phenols, alkylated phenols, polyphenols, monomeric and oligomeric phenols, creosol, thymol, alkoxyphenols, para-coumaryl alcohol, coniferyl alcohol, sinapyl alcohol, flavenols, catechins;

d) carboxylic acids such as formic acid, acetic acid, and phenolic acids such as ferric acid, benzoic acids, coumaric acid, cinnamic acid, abietic acid, oleic acid, linoleic acid, palmitic acid, stearic acid;

e) furans such as tetrahydrofuran (THF);

f) alkanes, alkenes, toluene, cumene.

The concentration of individual liquid organic compounds of the aqueous phase formed in the process in the aqueous phase entering the recovery system is often less than 2.0 wt.%, for example less than 1.0 wt.%.

However, in some embodiments of the present invention, for example, when other liquid aqueous phase organic compounds, such as alcohols or phenols, are added to the feed mixture, the concentration of a single aqueous phase liquid organic compound entering the recovery system can be up to 40 wt. .%, for example up to 30 wt.%, preferably up to 20 wt.%, for example in the range of 5-20 wt.%.

Liquid organic compounds of the aqueous phase may also contain emulsified droplets of the oil phase in accordance with the present invention.

Preferably, the recovery of aqueous phase liquid organics and homogeneous catalysts in the form of potassium and/or sodium from the aqueous phase comprises one or more of evaporation, distillation/fractionation, reverse osmosis, nanofiltration, ultrafiltration and pervaporation.

Often, the recovery of aqueous phase liquid organics and homogeneous catalysts in the form of potassium and/or sodium from the aqueous phase includes one or more evaporation and/or distillation steps, resulting in a first aqueous phase enriched in aqueous phase liquid organics and homogeneous catalysts in the form of potassium and/or sodium (concentrate), and a second aqueous phase stream depleted in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and/or sodium (distillate), where the amount of the resulting second aqueous phase is chosen to match the amount of water entering the high pressure processing system, for example contained in one or more carbonaceous feedstocks.

The aqueous phase entering the recovery system may, in accordance with an embodiment of the present invention, be filtered to remove suspended solids before entering one or more evaporation and/or distillation stages.

Often, the recovery system additionally includes one or more flashing stages.

The pH of the aqueous phase in the recovery system is preferably maintained in the alkaline range, eg in the pH range of 7-14, eg in the pH range of 9-14, preferably the pH is maintained in the range of 10-14, eg in the range of 10-13. Said alkaline maintenance often involves measuring and adjusting the pH by adding sodium hydroxide to the aqueous phase.

Advantageously, the evaporated vapor is brought into contact with the absorbent in the absorber prior to said condensation steps. Said absorber may contain an alkaline absorbent such as sodium hydroxide. The sodium hydroxide added to said absorber may be sodium hydroxide added to the aqueous phase to maintain the pH in the desired pH ranges during the recovery step and/or the effluent treatment step described above.

By maintaining the pH in the recovery system and/or the absorber stage within the above ranges, the concentration of phenols in the distillate fraction is reduced.

In a preferred embodiment, the recovery system comprises at least one evaporator, such as a falling film evaporator. Preferably, the evaporated vapor in said evaporation stage is condensed in at least two condensation stages having a decreasing temperature. Often, the evaporated vapor passes through the eliminator and/or coalescer prior to said condensation stage(s).

The recovery system according to a preferred embodiment of the present invention may comprise one or more distillation column(s) comprising a desorption and distillation section.

One or more carbonaceous feedstocks are selected from the following: biomass such as woody biomass and waste products such as wood chips, sawdust, timber harvested from thinning, road construction, bark, twigs, garden and park waste and weeds, energy crops, such as coppice, willow, miscanthus and reed arundo; agricultural residues and by-products such as grasses, straw, stalks, husks, corn cobs and hulls, eg wheat, rye, corn, rice, sunflower; empty fruit bundles from palm oil production, palm oil production effluents (ROME), residues from sugar production such as bagasse, vinasse, molasses, greenhouse waste; energy crops such as miscanthus, millet, sorghum, jatropha; aquatic biomass such as macroalgae, microalgae, cyanobacteria; bedding for livestock and compost, for example, the fibrous fraction of feed production; municipal and industrial waste streams, eg black liquor, paper sludge, off-grade pulp and paper fibers; residues and by-products from food production such as juice or wine production, vegetable oil production, sorted municipal solid waste, sorted household waste, restaurant waste, meat processing waste, sewage sludge, plastics, and combinations thereof.

By using such a separation method, as compared to previously known methods, the accumulation of undesirable components in the system can be avoided, which leads to undesirable effects on the process and process system.

It should be noted that the method is defined as comprising separating the product mixture into a gas phase, an oil phase (liquid hydrocarbon) and an aqueous phase containing liquid organic compounds of the aqueous phase, dissolved salts and optionally suspended particles. It means,

- 4 040945 that the phases contain mainly gas, liquid hydrocarbon and water, but also other components, while the subsequent separation process serves to further purify, in particular, the liquid hydrocarbon phase.

Brief description of the drawings

In the following, the invention will be described with reference to one embodiment shown in the drawings, where in FIG. 1 is a schematic diagram of an embodiment of a continuous high pressure process for converting carbonaceous materials into liquid organic compounds of a renewable oil phase;

in fig. 2 is a schematic diagram of a first embodiment of a continuous high pressure process for converting carbonaceous materials into liquid organic compounds of a renewable oil phase, including a system for recovering liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium according to the invention;

in fig. 3 is a schematic diagram of another embodiment of a continuous high pressure process for converting carbonaceous materials into liquid organic compounds of a renewable oil phase, including a system for recovering liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, and additionally including diverting an exhaust stream from an aqueous phase enriched in liquid organic compounds of the aqueous phase according to the invention;

in fig. 4 is a schematic diagram of an embodiment of the separation system of the invention;

in fig. 5 is a schematic drawing of a preferred embodiment of a 3-phase separator according to the invention;

in fig. 6 shows a schematic diagram of another embodiment of the separation system of the invention, further comprising a flash separator for removing low boiling compounds and water from the oil phase after the second phase separator;

in fig. 7 is a schematic diagram of a preferred embodiment of the separation system of the invention further comprising a flash separator for separating gas from the converted feed mixture prior to entering the first phase separator;

in fig. 8 shows a schematic diagram of a preferred embodiment of the separation system of the invention, including recirculation of the recovered flash ends and recirculation of the washing agent to the washing stage;

in fig. 9 is a schematic diagram of a preferred embodiment of a high pressure process for processing a feed stream containing carbonaceous material; the scheme includes a preferred separation system, including a recovery system for recovering liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium;

in fig. 10 is a schematic diagram of a preferred embodiment of the recovery system of the present invention, including an evaporation process;

in fig. 11 is a schematic diagram of another embodiment of a recovery system comprising two distillation columns for separating a process water stream;

in fig. 12 is a diagram of a preferred embodiment of a recovery unit comprising an evaporator and two distillation columns;

in fig. 13 is a schematic diagram of a preferred exhaust stream treatment system comprising a salt separation unit comprising a first filter and two fixed beds with a chloride selective ion exchange resin;

in fig. 14 is a schematic diagram of another preferred embodiment of a recovery system comprising a salt separation unit comprising a first filter and two fixed beds of chloride selective ion exchange resin, and wherein an additional effluent is drawn from the chloride-depleted water stream exiting the salt separation unit.

Description of the Preferred Embodiment

In FIG. 1 shows an embodiment of a continuous high pressure production process for converting carbonaceous materials such as biomass to renewable oil, comprising the following.

1. The stage of preparation of the raw mixture.

2. Stage conversion, including stages

a) pressurization;

b) heating;

c) reactions.

3. Cooling and pressure reduction.

4. Separation.

5. Quality improvement.

- 5 040945

1. Preparation of raw materials.

The first step (1) of the process consists in preparing the feed mixture in the form of a pumpable suspension of carbonaceous material. It typically includes size reduction and suspending means, such as dispersing the organic material with other ingredients such as water, catalysts, and other additives such as organic compounds in the feed mixture.

The carbonaceous material of the present invention may be in solid form or may have the appearance of a solid, but may also be in the form of a slurry or liquid. In addition, carbonaceous material(s) may be contained in one or more incoming streams.

Non-limiting examples of carbonaceous feedstocks of the present invention include the following: biomass such as woody biomass and waste products such as wood chips, sawdust, timber harvested from thinning, road construction, bark, twigs, garden and park waste and weeds, energy crops , such as coppice, willow, miscanthus and reed arundo; agricultural residues and by-products such as grasses, straw, stalks, husks, corn cobs and hulls, eg wheat, rye, corn, rice, sunflower; empty fruit bundles from palm oil production, palm oil production effluents (ROME), residues from sugar production such as bagasse, vinasse, molasses, greenhouse waste; energy crops such as miscanthus, millet, sorghum, jatropha; aquatic biomass such as macroalgae, microalgae, cyanobacteria; bedding for livestock and compost, for example, the fibrous fraction of feed production; municipal and industrial waste streams, eg black liquor, paper sludge, off-grade paper fibers; residues and by-products from food production, such as juice or wine production; vegetable oil production, sorted municipal solid waste, sorted household waste, restaurant waste, meat processing waste, sewage sludge, and combinations thereof.

Many of the carbonaceous materials of the present invention relate to lignocellulosic materials such as woody biomass and agricultural residues. Such carbonaceous materials typically contain lignin, cellulose and hemicellulose.

An embodiment of the present invention includes a carbonaceous material having a lignin content in the range of 1.0-60% by weight, such as a lignin content in the range of 10-55% by weight. Preferably the content of lignin in the carbonaceous material is in the range of 15-40 wt.%, for example 20-40 wt.%.

The content of cellulose in the carbonaceous material is preferably in the range of 10-60 wt.%, for example, the content of cellulose is in the range of 15-45 wt.%. Preferably, the content of cellulose in the carbonaceous material is in the range of 20-40 wt.%, for example 30-40 wt.%.

The content of hemicellulose in the carbonaceous material is preferably in the range of 10-60% by weight, for example the content of hemicellulose is in the range of 15-45% by weight. Preferably, the content of cellulose in the carbonaceous material is in the range of 20-40 wt.%, for example 30-40 wt.%.

Depending on the particular organic matter to be transformed and how it is produced, the size reduction may be carried out in one or more steps, for example, the carbonaceous material may be processed as is and subsequently blended with other ingredients in the same step, or it may be preliminarily crushed to a size suitable for further processing and size reduction at the mixing stage. Often, the carbonaceous material is ground to a particle size of less than 15 mm, for example to a particle size of less than 10 mm, in a pre-grinding step; preferably to a particle size of less than 5 mm, for example less than 3 mm.

Pre-grinding in accordance with an embodiment of the present invention may be carried out using a shredder, cutting mill, hammer mill, runner mill, impeller grinder, or a combination thereof.

Preferably, the pre-grinding step may further include means for removing impurities such as metals, stones, dirt like sand, and/or for separating off-grade fibers from the carbonaceous material of the specified maximum particle size. Such means may include magnetic separation, washing, density separation such as flotation, vibrating tables, acoustic separators, screening, and combinations thereof. These agents may be present before the pre-grinding step and/or after the pre-grinding step.

The carbonaceous material is subsequently mixed with the other ingredients of the raw mix. Other ingredients may include the following.

1. Recycle oil (hydrocarbons) obtained in the process, or oil fraction (hydrocarbon obtained in the process) preferably in a mass ratio to dry ashless organic matter in the range from 0.5 to 1.5, for example, in a ratio of 0.8 up to 1.2; Recycle oil may contain

- 6 040945 harvest phenols, alkylated phenols, polyphenols, monomeric and oligomeric phenols, creosol, thymol, alkoxyphenols, para-coumaryl alcohol, coniferyl alcohol, sinapyl alcohol, flavenols, catechins.

2. Recycled aqueous phase concentrate from the process, containing recovered homogeneous catalyst and water-soluble organics, such as one or more components selected from the following:

a) ketones such as acetone, propanones, butanones, pentanones, pentenones, cyclopentanones such as 2,5-dimethylcyclopentanone, cyclopentenones, hexanones and cyclohexanones, for example 3methylhexanone, quinones, etc.;

b) alcohols and polyalcohols such as methanol, ethanol, propanols (including isopropanol), butanols, pentanols, hexanols, heptanols, octanols such as 2-butyl-1-octanol, hydroquinones, benzenediols, etc.;

c) phenols, alkylated phenols, polyphenols, monomeric and oligomeric phenols, creosol, thymol, alkoxyphenols, para-coumaryl alcohol, coniferyl alcohol, sinapyl alcohol, flavenols, catechins;

d) carboxylic acids such as formic acid, acetic acid, and phenolic acids such as ferric acid, benzoic acids, coumaric acid, cinnamic acid, abietic acid, oleic acid, linoleic acid, palmitic acid, stearic acid;

e) furans such as THF etc.;

f) alkanes, alkenes, toluene, cumene, xylene, etc.; and combinations of this.

Typically, water-soluble organic compounds form a complex mixture of the above components, and the raw mixture may contain such water-soluble organic compounds at a concentration of from about 1 to about 10 wt.%, for example, in the range from about 2 to about 5 wt.%.

3. Additional homogeneous catalyst in the form of potassium carbonate and/or potassium hydroxide and/or potassium acetate; preferably added as an aqueous solution and added in such an amount that the total concentration of potassium in the resulting raw mixture is at least 0.5 wt.%, such as the concentration in the raw mixture of at least 1.0 wt.%; preferably the concentration of potassium is at least 1.5 wt.%, for example at least 2.0 wt.%.

4. Additive base for pH adjustment. Preferably, sodium hydroxide is added to the feed mixture in such an amount that the pH measured in the recycled aqueous phase is above 7 and is preferably in the range of 8.0 to 12.0, for example in the range of 8.0 to 10.0.

All of the ingredients of claims 1 to 4 are preferably in liquid form and may advantageously be premixed and optionally preheated prior to mixing with organic matter to obtain said raw mixture. Pre-mixing and/or pre-heating can reduce the loading time and heating time required in the mixer.

The carbonaceous material and other ingredients are mixed in such a way that a homogeneous suspension or paste is formed. Said mixer may be an agitated vessel equipped with means for efficient mixing, dispersion and homogenization of viscous materials, such as a planetary mixer, a kneader or a Banbury mixer.

The mixer is preferably further provided with means for preheating said raw mixture to a temperature in the range of about 80 to about 250°C, preferably in the range of about 130 to about 220°C, and more preferably in the range of about 150 to about 200°C, for example in the range from about 160 to about 180°C, at a pressure sufficient to prevent boiling, for example a pressure in the range of 1-30 bar (0.1-3 MPa), preferably in the range of 4-20 bar (0.4-2 MPa ), for example in the range of 5-10 bar (0.5-1.0 MPa).

Heating the feed mixture to temperatures in the above ranges results in softening and/or at least partial dissolution of carbonaceous substances, which facilitates grinding and homogenization of the feed mixture. Preferred means for heating said raw mixture during the preparation of the present invention include a heating jacket. In a preferred embodiment, the heat for preheating said feed mixture is obtained by cooling the converted carbonaceous material containing the liquid hydrocarbon product. This makes it possible to further improve the energy efficiency of the process. The mixer may be further equipped with a recirculation circuit, whereby material is withdrawn from said mixer and at least partially recirculated in an internal or external circuit, and re-introduced into said mixer in order to control the characteristics of the raw mixture, for example, rheological properties such as viscosity and/or particle size, to a predetermined level. The outer circuit may further comprise one or more size reduction and/or homogenization devices such as a macerator and/or colloid mill and/or cone mill, or a combination thereof, in series and/or parallel arrangement.

It is preferable that the carbonaceous material be fed into the mixer gradually rather than all at once to control the viscosity of the raw mixture and to keep the raw mixture pumpable.

- 7 040945 when reducing the size and homogenization. Viscosity control can be performed by measuring the energy consumption of the mixer and/or colloid mill, and by adding organic matter to the raw mixture in accordance with a predetermined energy consumption. In addition, it is preferable not to empty the mixer completely between batches, since the prepared raw mix acts as a texturizing agent for the next batch and thereby helps to homogenize the next batch, making it more pumpable and thus the carbonaceous material can be added more quickly.

Other preferred means for thoroughly mixing and homogenizing the ingredients in the raw mix include inline mixers. Such inline mixers also allow cutting and/or shearing and/or self-cleaning action. A preferred embodiment of such an inline device includes one or more extruders.

The feed mixture from the feed mixture mixing step may be fed into a holding tank before entering the pressurization step of the process. Said mixing tank may be provided with means for mixing said feed mixture in a storage tank and/or circulation means for circulating said feed mixture around said storage tank, while keeping the feed mixture in a shear thinned and more easily pumpable state. Optionally, the raw mixture may be expanded prior to entering the storage tank, whereby the raw mixture may be further crushed and homogenized.

Generally, the dry weight content of the carbonaceous material in the raw material mixture of the present invention is in the range of 10-40 wt.%, preferably in the range of 15-35 wt.%, and more preferably in the range of 20-35 wt.%.

The method according to the present invention requires the presence of water in said raw mixture. Typically, the water content of said raw material mixture is at least 30% by weight, in the range of 30-80% by weight, and preferably in the range of 40-60% by weight.

2. Conversion.

The second stage, conversion, includes a pressurization step (2a) in which the pressure of the feed mixture is increased by means of pumping means to a pressure of at least 150 bar (15 MPa) to about 450 bar (45 MPa), for example from a pressure of at least 180 bar (18 MPa) to 400 bar (40 MPa); preferably, the pressure of the raw mixture is increased by means of pumping means to a pressure above the critical point of water, such as a pressure of at least 250 bar (25 MPa); more preferably, the pressure of the raw mixture is increased by pump means to a pressure of at least 300 bar (30 MPa), such as at least 320 bar (32 MPa). A particularly preferred embodiment of the present invention is one in which the pressure of the raw mixture downstream of the pump means is 320-380 bar (32-38 MPa). According to the present invention, said pressurization to the desired reaction pressure is essentially carried out before heating from the inlet temperature of the raw mix preparation stage to the reaction temperature.

Many embodiments of the present invention relate to the processing of raw mixtures with a high content of carbonaceous material, as described above. Such raw mixes typically have a density in the range of 1050-1200 kg/m ³ and typically behave like a homogeneous pseudo-plastic paste rather than a suspension of discrete particles (liquid). The viscosity of such pastes can vary widely with shear rate due to pseudoplastic behavior (shear thinning) and can range from 10 ³ to 10 ⁷ cP (1-10,000 Pa-s) depending on the particular shear rate and the carbonaceous material being processed.

One aspect of the present invention relates to a pressurization system for pressurizing such highly viscous pseudoplastic raw mixes. In accordance with a preferred embodiment of the present invention, the pressure boosting system comprises two or more pressure boosters, each containing cylinders with a piston provided with actuating means for applying and/or receiving force on the piston. Preferred driving means for the pistons in the cylinders of the present invention include hydraulic driving means.

The pressure reducing device of the present invention is generally designed for low stroke speeds (high displacement), which allows actuated valves rather than check valves to be used to fill and empty the cylinders. Preferred actuated valves of the present invention are gate valves and ball valves, or a combination thereof.

The stroke speed of the pistons in accordance with an embodiment of the present invention may be from about 1 stroke/min to about 150 strokes/min, such as from about 5 strokes/min to about 100 strokes/min. Preferably, the stroke rate of the pistons is from about 10 to about 80 strokes/min, for example, the stroke rate of the pistons is in the range from about 20 to about 60 strokes/min. In addition to the possibility of using actuated valves, the low piston stroke speed reduces wear on pistons, seals and valve seats.

Inlet temperature at pressurization is typically in the range of about 10 to

- 8 040945 about 250°C, for example from about 20 to about 220°C; preferably, the inlet temperature to the pressurizing cylinders is typically in the range of from about 50 to about 210°C, for example from about 80 to about 200°C; even more preferably, the inlet temperature to the pressurizing cylinders is typically in the range of about 100 to about 180°C, eg, about 120 to about 170°C.

For applications of the present invention, where the temperature exceeds about 120°C, for example about 140°C, the cylinders can be additionally equipped with means for cooling the piston seals to withstand operating conditions.

In a preferred embodiment, the pressure energy is recovered in the depressurization step described below as step 6. Depressurization and transfer to an energy store takes place, where the energy absorbed by the pressure reducer is transferred to the accumulator for later use, e.g., in the pressurization step. The result is a very energy efficient high pressure process.

The pressurized feed mixture is further heated (2b) to a reaction temperature in the range from about 300 to about 450°C, for example, to a temperature in the range from about 330 to about 430°C; preferably, the pressurized feed mixture is further heated to a reaction temperature in the range of from about 350 to about 425°C, for example, to a temperature in the range from about 390 to about 420°C, for example, in the range from 400 to 415°C.

In accordance with one aspect of the present invention, the heating of the raw mixture is carried out by indirect heat exchange with high pressure water as a heat transfer medium between the cooling and heating stages. By using such a heat transfer medium, both the raw material mixture and the product mixture can flow inside the tubes, which makes cleaning easier.

By said heat recovery, it is achieved that the process becomes very energy efficient, since most of the required heat is recovered. In many embodiments of the present invention, at least 40% of the energy required to heat the feed mixture to the desired reaction temperature is recovered, for example, at least 50% of the energy required to heat the feed mixture to the desired reaction temperature is recovered. Preferably at least 60% of the energy required to heat the feed mixture to the desired reaction temperature is recovered, eg at least 70% of the required energy is recovered.

After heating to the reaction temperature, said pressurized and heated feed mixture is maintained at the required pressure and temperature in the reaction zone (2c) for a predetermined time to convert the carbonaceous material(s). The feed characteristics and/or combination of pressure and temperature of the present invention generally provide shorter reaction times and/or more liquid hydrocarbon product reacted than in the prior art while maintaining the yield and/or quality of the desired product. The specified residence time in said reaction zone according to an embodiment of the present invention may be in the range of 1-60 minutes, such as 2-45 minutes, preferably said specified residence time in said reaction zone is in the range of 3-30 minutes, such as in the range of 3 -25 minutes, more preferably in the range of 4-20 minutes, such as 5-15 minutes.

3. Cooling and expansion.

A mixture of products containing liquid hydrocarbon product, water with liquid organic compounds of the aqueous phase and dissolved salts, gas containing carbon dioxide, hydrogen and methane, as well as suspended particles from the specified converted carbon-containing material, is then cooled (3) to a temperature in the range of 70 up to 250°C, for example, in the range from 120 to 220°C; preferably up to a temperature in the range from 130 to 200°C, for example, in the range from 140 to 180°C.

A preferred embodiment of the cooling step of the present invention is that said heat exchange is carried out by indirect heat transfer using high pressure water as the heat transfer medium, as described in the conversion section. Through the use of such indirect heat transfer through the heat transfer medium, both the raw material mixture and the product mixture can flow inside the tubes, thereby allowing easier cleaning. The heat transfer medium may optionally be further heated and/or further cooled to provide additional heating and cooling control and flexibility. Said heat transfer medium can also be used to transfer heat to/from other process steps of the process, such as, for example, the preparation of raw materials (1) and/or the improvement section of the process of the present invention.

The cooled product mixture then enters the pressure reducing device (3) where the pressure is reduced from the conversion pressure to a pressure of less than 200 bar (20 MPa), for example, to a pressure of less than 120 bar (12 MPa). Preferably the pressure is reduced to less than 90 bar (9 MPa), for example less than 80 bar (8 MPa). More preferably, the pressure is reduced to less than 50 bar (5 MPa), for example to a pressure in the range of 10-40 bar (1-4 MPa).

Suitable pressure reducing devices include pressure reducing devices comprising a series of tubular elements in series and/or parallel arrangement with a length and internal

- 9 040945 cross section, adapted to reduce the pressure to the desired level, and pressure reducing devices containing pressure reducing pump units.

In a preferred embodiment, the cooled product mixture enters the pressure reduction device, wherein the pressure reduction unit comprises at least one inlet and outlet, the pressure reduction unit is configured to receive pressurized fluid at the inlet process pressure, is configured to isolate the resulting pressurized fluid from the upstream process and from the outlet, and is configured to reduce the pressure of the fluid to a lower predetermined level, and is also configured to divert fluid through the outlet, while maintaining isolation with respect to the process located upstream flow.

Typically, the pressure reducing unit comprises an inlet actuated valve and an outlet actuated valve, and between the inlet valve and the outlet valve, a pressure booster. In addition, the pressure reducing unit according to an embodiment of the present invention includes means for measuring pressure before the inlet valve, between the inlet valve and the outlet valve, and after the outlet valve.

The pressure reduction unit of the present invention may also comprise a pumping unit having a cylinder and a piston, as well as means for driving the piston within the cylinder. Preferably, the pressure reducing unit also includes a position indicator indicating the cycle position of the pressure reducing device and configured to provide a control signal to open or close at least one valve in the pressure reducing system.

A preferred embodiment of the pressure reducing device of the present invention is that the pressure reducing pump is connected to an additional pump which controls the pressurization of the energy store. For example, the pressure reducing device further comprises an energy storage device, wherein the pressure pump is operatively connected to the storage device, and the energy absorbed by the pump is converted and transferred to the pressure pump.

In a preferred embodiment, the energy store drives a pressure pump capable of pressurizing the raw mix in the pressurizing step (step 2 above) of the high pressure process. In one embodiment of the present invention, this is done by a low pressure turbine connected to a generator generating electricity, and the generated electricity reduces the energy required to drive the pressure pump during the pressurization stage.

The pressure reducing device of the present invention is generally designed for low stroke speeds (high displacement), which allows actuated valves rather than check valves to be used to fill and empty the cylinders. Preferred actuated valves of the present invention are gate valves and ball valves, or a combination thereof.

The stroke speed of the pistons in accordance with an embodiment of the present invention may be from about 1 stroke per minute to about 150 strokes per minute, for example, from about 5 strokes per minute to about 100 strokes per minute. Preferably, the piston stroke rate is from about 10 to about 80 strokes per minute, for example, the piston stroke rate is in the range from about 20 to about 60 strokes per minute. In addition to the possibility of using actuated valves, the low piston stroke speed reduces wear on pistons, seals and valve seats.

The inlet temperature to the pressure reducing device is typically in the range of from about 10 to about 250°C, for example, from about 20 to about 220°C; preferably, the inlet temperature to the pressurizing cylinders is typically in the range of from about 50 to about 210°C, for example from about 80 to about 200°C; even more preferably, the inlet temperature to the pressurizing cylinders is typically in the range of about 100 to about 180°C, eg, about 120 to about 170°C.

For applications of the present invention, where the temperature exceeds about 120°C, for example about 140°C, the cylinders can be additionally equipped with means for cooling the piston seals to withstand operating conditions.

4. Separation.

The converted feed mixture is further separated (4) into at least a gas phase containing carbon dioxide, hydrogen, carbon monoxide, methane and other short chain hydrocarbons (C ₂ -C ₄ ), alcohols and ketones, a crude oil phase, an aqueous phase with liquid aqueous phase organic compounds, as well as dissolved salts and, in some cases, suspended particles such as inorganic substances and/or coal and/or unconverted carbonaceous material, depending on the particular carbonaceous material being processed and the particular processing conditions. Dissolved salts and inorganic substances may include metal or alkali or alkaline earth metals such as potassium, sodium, chlorides, sulfate, carbonate and bicarbonate, aluminum, calcium, magnesium, sodium and potassium, silicon dioxide, iron, cobalt, nickel, phosphorus. Inorganic materials are formed from carbonaceous raw materials such as biomass and/or from homogeneous catalyst(s) used in the high pressure production process and/or as a result of fouling in the high pressure production process.

For some carbonaceous materials having a high inorganic content, the product stream, partially cooled and partially reduced pressure, can be filtered to remove suspended solids before entering further separation (4).

According to a preferred embodiment, the separation is carried out by first separating the individual phases in a phase separator, such as a 3-phase separator, and subsequent purification of the separated oil phase, such as reducing the concentration of contaminants such as water and/or inorganic substances, for example by adding one or more flushing agents and/or viscosity reducers and/or density reducing agents, and by separating the oil phase from one or more flushing agents and/or viscosity reducers and/or density reducing agents in a 3-phase separator.

The aqueous phase from the first separator typically contains homogeneous catalyst(s) such as potassium and sodium, as well as liquid organic compounds of the aqueous phase.

5. Quality improvement.

The renewable crude oil may be further sent to an upgrading process (5) where it is pressurized to a pressure in the range of about 20 bar to about 200 bar (2-20 MPa), for example to a pressure in the range of 50-120 bar (5 -12 MPa) before heating to a temperature in the range of 300-400°C in one or more stages and contacting with hydrogen and heterogeneous catalyst (s) contained in one or more reaction zones, and ultimately separated into fractions with different boiling points .

In FIG. 2 is a schematic diagram of an embodiment of a continuous high pressure process for converting carbonaceous materials into liquid organic compounds of a renewable oil phase, further including a system for recovering liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium.

Liquid organic compounds of the aqueous phase often contain a complex mixture and usually contain one or more compounds selected from one or more of the following groups:

a) ketones such as acetone, propanones, butanones, pentanones, pentenones, cyclopentanones such as 2,5-dimethylcyclopentanone, cyclopentenones, hexanones and cyclohexanones, for example 3methylhexanone, quinones;

b) alcohols and polyalcohols such as methanol, ethanol, propanols, butanols, pentanols, hexanols, heptanols, octanols such as 2-butyl-1-octanol, hydroquinones, benzene diols;

c) phenols, alkylated phenols, polyphenols, monomeric and oligomeric phenols, creosol, thymol, alkoxyphenols, para-coumaryl alcohol, coniferyl alcohol, sinapyl alcohol, flavenols, catechins;

d) carboxylic acids such as formic acid, acetic acid, and phenolic acids such as ferric acid, benzoic acids, coumaric acid, cinnamic acid, abietic acid, oleic acid, linoleic acid, palmitic acid, stearic acid;

e) furans such as tetrahydrofuran (THF);

f) alkanes, alkenes, benzene, toluene, cumene, xylene.

The aqueous phase from separation step (4) according to a preferred embodiment of the present invention is fed to a recovery system for recovering liquid organic compounds of the aqueous phase and/or homogeneous catalysts in the form of potassium and sodium salts.

Many preferred embodiments of the continuous high pressure processing of carbonaceous material to hydrocarbons of the present invention include a recovery step to recover homogeneous catalyst(s) and/or aqueous phase liquid organic compounds from the aqueous phase from separation step (4). This results in an aqueous phase depleted in liquid organic compounds and homogeneous catalysts in the form of potassium and sodium, and an aqueous phase enriched in liquid organic compounds and homogeneous catalysts in the form of potassium and sodium. The liquid phase enriched in aqueous phase liquid compounds and homogeneous catalysts in the form of potassium and sodium is in a preferred embodiment at least partially recycled and introduced into the feed preparation step as shown in the figure. Thus, the overall oil yield and the energy efficiency of the process are increased and the economy of the process is greatly improved by said recovery and recycling of homogeneous catalysts.

In a preferred embodiment of the present invention, the recovery system includes one or more of evaporation, distillation/fractionation, reverse osmosis, nanofiltration, ultrafiltration, pervaporation, activated carbon, a biological wastewater treatment step, and a combination thereof.

In a preferred embodiment, the recovery system (6) includes an evaporation stage and/or one or more distillation stages, where heat for evaporation and/or distillation is at least partially provided by heat transfer from a high pressure water cooler through a heat transfer medium such as hot oil. or steam, while overall heat recovery and energy efficiency are increased.

The liquid organic compounds of the oil phase are also, in a preferred embodiment, recycled and introduced into the feed preparation step, as also shown in FIG. 2.

In FIG. 3 is a schematic diagram of a preferred embodiment of a continuous high pressure process for converting carbonaceous materials into liquid organic compounds of a renewable oil phase, comprising a system for recovering liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, and further comprising diverting an effluent stream from the aqueous phase. enriched with liquid organic compounds of the aqueous phase and homogeneous catalysts containing potassium and/or sodium before being introduced to the stage of preparation of raw materials.

The aqueous phase from the separation system contains liquid organic compounds of the aqueous phase and dissolved homogeneous catalyst, and may also contain suspended particles and other dissolved salts. The aqueous phase, in accordance with a preferred embodiment of the invention, may be filtered before entering the recovery unit to reduce the amount of suspended particles. In this way, fouling of the recovery system can be reduced and cleaning and maintenance intervals are extended, which increases the overall availability of the process.

An additive base, such as sodium hydroxide, may be added to the process water prior to entering the recovery system to maintain the pH of the process water in the recovery system in the range of 7 to 14, for example, in the range of 8.5 to 14; preferably in the range of 9 to 14, for example in the range of 10 to 14; even more preferably, the pH of the process water entering the recovery system is maintained in the range of 10-13 by measuring the pH and adding base to the process water prior to entering the recovery system. As a result, the volatility of aqueous phase liquid organic compounds such as phenols is reduced and thus more retained in the aqueous phase enriched in aqueous phase liquid organic compounds (concentrate) when the evaporation and/or distillation methods of the present invention are applied. Thus, further processing of the depleted aqueous phase is simplified and may even be eliminated in some embodiments of the present invention, for example, an aqueous phase depleted in liquid organic compounds of the aqueous phase can be treated sufficiently for direct discharge.

However, while impurity elements such as most divalent ions such as calcium and metals have limited solubility in the aqueous phase and will be removed as suspended solids in the separation and filtration system(s), it has been found that dissolved salts, such as chloride will accumulate without diverting the exhaust stream. Therefore, according to a preferred embodiment, an effluent is drawn from an aqueous phase rich in liquid organics of the aqueous phase and homogeneous catalysts, as shown in FIG. 3.

The minimum amount of exhaust flow required is determined by the concentration of chloride in the carbonaceous material, i.e. the amount of chloride supplied with the carbonaceous material must be equal to the amount of chloride removed with the outlet stream.

According to a preferred embodiment of the present invention, the mass ratio of the outlet stream withdrawn from the aqueous phase stream enriched in liquid organic compounds and homogeneous catalyst to the total aqueous phase stream fed to said recovery system is in the range of 0.01 to 0.5, for example in the range from 0.02 to 0.4; preferably, the mass ratio of said discharged outlet stream to the total aqueous phase stream entering the recovery system is in the range of 0.03-0.25, for example in the range of 0.04-0.15.

In FIG. 4 shows a schematic diagram of a first embodiment of the separation system of the present invention. The conversion product is cooled to a temperature in the range from 50 to 250°C, for example to a temperature in the range from 60 to 220°C, preferably to a temperature in the range from 120 to 180°C and most preferably to a temperature in the range from 130 to 170°C , and is depressurized to a pressure in the range of 10-150 bar (1-15 MPa), for example to a pressure in the range of 10-100 bar (1-10 MPa), preferably the conversion product is depressurized to a pressure in the range of 10-74 bar (1-7.4 MPa), for example up to a pressure in the range of 15-50 bar (1.5-5 MPa), even more preferably up to a pressure in the range of 20-50 bar (2-5 MPa).

The product stream, partially cooled and partially reduced pressure, from the conversion stage is fed to the first phase separator, where the conversion product is separated under pressure into a gas phase, an oil phase and an aqueous phase and, optionally, a solid phase, depending on the specific converted carbonaceous material and specific operating conditions of the conversion process.

In accordance with many embodiments of the present invention, the first separator is a gravimetric phase separator, as further shown in FIG. 5. The phase separator may according to the present invention be positioned horizontally or vertically, however in many preferred applications according to the present invention the first three phase separator is positioned horizontally. With a horizontal arrangement of the phase separator

- 12 040945 get a large interface between gas and liquid, thereby achieving a minimum collision of gas bubbles moving up, and liquid drops moving down. In this way a more efficient separation is achieved, for example the separation efficiency can be increased and/or a shorter residence time can be used.

The first phase separator comprises an inlet for introducing said product mixture and outlets for withdrawing a gas phase, an oil phase (liquid hydrocarbon), an aqueous phase and, optionally, a solid phase.

The operating temperature of the first phase separator in the preferred embodiment is chosen so as to obtain a dynamic viscosity of the liquid hydrocarbon product in the range of from about 0.1 centipoise to about 30 centipoise (0.001-0.03 Pa-s) during said additional separation, for example, in the range of from about 1 centipoise to about 20 centipoise (0.001-0.020 Pa-s) during said additional separation, preferably the separation temperature is chosen so as to obtain a dynamic viscosity in the range from about 1 centipoise to about 20 centipoise (0.0010, 02 Pa-s), for example in the range of 5-15 cP (0.005-0.015 Pa-s).

The operating temperature of the first phase separator may according to an embodiment of the present invention be in the range of 50 to 250°C, for example in the range of 80 to 200°C, preferably the operating temperature in the first phase separator is in the range of 120 to 180°C, for example in the range from 130 to 170°C. By maintaining the operating temperature of the first separator in the specified range, it is obtained that the dynamic viscosity of the liquid hydrocarbon product (oil phase) is maintained in the above range, thereby increasing the separation efficiency of water and/or particles contained in the oil phase.

In addition, it was found that the oil phase may contain a large amount of organic compounds that have a melting point in the range from about 100 to 120°C. Such organic compounds may contain high molecular weight compounds, such as organic resins and/or asphaltene-like compounds, which can solidify on inorganic particles in the oil and/or stabilize water droplets in the oil phase. Such stabilization may result from an interfacial film consisting of surface-active high molecular weight polar solids coating small water droplets, and this interfacial film provides a barrier to coalescence of droplets at too low a temperature. By maintaining the operating temperature of the separator sufficiently high (eg, above the melting point of such compounds), separation efficiency can be improved by the present invention.

The operating pressure of the first phase separator in accordance with the present invention is usually chosen above the boiling pressure of the liquid phase, so that the liquid phases are essentially maintained in their liquid state at the prevailing separation temperature. Therefore, in many embodiments of the present invention, the operating pressure of the first phase separator is at least 5 bar (0.5 MPa), for example, the operating pressure is at least 10 bar (1 MPa).

However, it has been found that operating at a higher pressure improves separation, as will be further illustrated in the separation examples. Therefore, in a preferred embodiment of the present invention, the operating pressure of said first phase separator is in the range of 10-150 bar (1-15 MPa), for example in the range of 10-100 bar (1-10 MPa), preferably the pressure in the first separator is in the range 10-74 bar (1-7.4 MPa), for example in the range of 15-50 bar (1.5-5 MPa), and even more preferably in the range of 20-40 bar (2-4 MPa).

Many aspects of the present invention relate to the use of one or more phase separators, wherein the residence time in each of the phase separators is in the range of 160 minutes, for example, in the range of 1-30 minutes, preferably the residence time in each of the phase separators is in the range of 2- 20 minutes.

According to the present invention, the partially dehydrated and partially deashed oil phase is withdrawn from the first separator and subjected to a further purification process as shown in the figure.

In an aspect of the present invention, a portion of the oil phase from the first separator is withdrawn prior to further purification and recycled to the high pressure process feedstock preparation stage. Thereby, the dimensions of the second phase separator are reduced.

According to preferred embodiments of the present invention, the oil refining process includes mixing an oil phase with one or more flushing agents and then supplying the mixed oil phase and flushing agent to a second phase separator where it is separated into a phase containing at least one flushing agent and having an increased content of water and/or inorganic substances, and an oil phase having a reduced content of inorganic substances and/or water, and optionally a gas phase.

The operating pressure of the second separator according to the preferred embodiments of the present invention is in the range of 5-100 bar (0.5-10 MPa), preferably the pressure in the second separator is in the range of 10-74 bar (1-7.4 MPa), for example in the range 15-50 bar (1.5-13 040945

MPa), and even more preferably in the range of 20-40 bar (2-4 MPa).

The operating temperature of the second phase separator may according to an embodiment of the present invention be in the range of 50 to 250°C, for example in the range of 80 to 200°C, preferably the second phase separator is operated at a temperature in the range of 120 to 180°C, for example in the range from 130 to 170°C. By maintaining the separation operating temperature in the above range, it is obtained that the dynamic viscosity of the liquid hydrocarbon product (oil phase) is maintained in the above range, thereby increasing the separation efficiency of water and/or particles contained in the oil phase.

In many aspects of the present invention, the flushing agent may contain a viscosity and/or density reducer. The viscosity and/or density reducer may be an organic solvent having a boiling point below 150°C, eg below 140°C, preferably below 130°C, eg below 100°C.

Suitable viscosity and/or density reducers according to the present invention often contain at least one ketone selected from such as methyl ethyl ketone (MEK, 2-butanone), acetone, propanones, butanones, pentanones, pentenones, cyclopentanones such as 2,5dimethylcyclopentanone , cyclopentenones, hexanones, cyclohexanones such as 3-methylhexanones, 2heptanone and/or combinations thereof. Particularly preferred viscosity reducers of the present invention are methyl ethyl ketones and/or a low-boiling oil fraction from a converted feed mixture containing carbonaceous material.

The weight ratio of the viscosity and/or density reducer added to a certain amount of oil may be in the range of 0.01-2, for example in the range of 0.2-1, for example in the range of 0.2 to 0.5.

The viscosity reducer reduces the viscosity of the oil phase and may also reduce the density of the oil phase. In addition, the viscosity reducer can improve the dissolution of organic particles and/or improve the hydrophobicity of the oil phase. In this way, the separation efficiency is increased and/or the required separation time can be reduced.

In an aspect of the present invention, one or more flushing agents may contain one or more demulsifiers selected from xylenes, phenol-formaldehyde resin, n-propanol, heavy and light aromatic naphtha, ethylbenzene, 1,2,4-trimethylbenzene, 1,3,5-trimethylbenzene , 1,2,3 trimethylbenzene, glutaraldehyde, water, toluene, 2-butanone, ethyl acetate, 1-propyl acetate, or combinations thereof.

The demulsifier and/or mixture thereof require a concentration in the range of 10-50000 ppm by weight, for example in the range of 100-20000 ppm by weight, preferably in the range of 800-15000 ppm, for example in the range of 1000-10000 ppm

In many embodiments of the present invention, at least one of the flushing agents contains water. In yet another preferred embodiment of the present invention, at least one acidifying agent is added to at least one washing agent containing water. Suitable acidifying agents of the present invention include acetic acid and/or citric acid. Typically, said acidifying agent is added in such an amount that the pH of the separated pressure wash from the second separator is in the range of about 2 to about 7, for example in the range of about 2.5 to about 6.5; preferably, the pH of the separated flushing agent is in the range of from about 3 to about 6, for example, in the range from about 3 to about 5. By lowering the pH to these ranges according to the present invention, it is found that compounds such as potassium and sodium, which can be bound with acidic groups of oils, such as soaps, dissolve. In addition, the solubility of metals is also increased by lowering the pH. In addition, at too low a pH, it has been found that stable emulsions can form.

In a particularly preferred embodiment of the present invention, the acidifying agent comprises a pressurized gas obtained from a carbonaceous material conversion process. The process gas typically contains carbon dioxide as well as some light hydrocarbon gases such as methane, ethane, ethene, propane, propene, butane, butene, pentane, as further shown in Example 1. Typically, said process gas is withdrawn from the first separator, as shown in the figure and mixed with the flushing agent(s) in an inline mixer such as a static mixer before being introduced into the second phase separator. At operating pressures of the second phase separator of the present invention, CO ₂ dissolves in the water phase and forms carbonic acid, whereby the water is acidified to a pH in the range of 2.5 to 4. In addition, under operating conditions, the light hydrocarbon gases mentioned above can be dissolved in the oil phase, resulting in reduced oil viscosity and/or reduced oil phase density and/or improved oil phase hydrophobicity. Thereby, the separation efficiency is increased, as will be shown in the following examples. Another advantage of using process gas as an acidifying agent is that it easily separates from the oil product and/or flushing agent when the pressure is reduced to ambient values, which facilitates further processing.

- 14 040945 of these streams.

In FIG. 5 is a schematic drawing of a preferred embodiment of a 3-phase separator according to the invention. The product mixture preferably enters the phase separator through a product inlet (1) located in the free space above the liquid level at one end of the separator. The product mixture inlet is preferably provided with a flow diverter or distributor (2), such as a diffuser, to reduce fluid momentum and separate gas from liquids, resulting in a more efficient separation of gas and liquid. In other aspects of the present invention, the product inlet may contain or additionally contain cyclones or groups of cyclones (2).

In an alternative preferred embodiment, the separator may include a flash separator/degasser where the gas is separated from the product mixture and the liquid product mixture is introduced into the separator through a dip below the water phase (not shown in the figure).

In many preferred embodiments, the three-phase separator is further equipped with flow distribution means, wave and foam breaking devices such as perforated baffles (3), lamellar plates (4) or flow stabilization mesh as shown in the figure. The three-phase separator according to the present invention may, in additional aspects, also include coalescing devices (5) such as mesh, lamellar plates and/or electrocoalescing devices to speed up the coalescence process, thereby achieving a more efficient phase separation.

The three-phase separator according to the embodiments of the present invention usually also contains one or more chopping grids (6) for separating the liquid phases. Overflow of the oil phase is often observed, as shown in the figure.

The gas is typically vented from the outlet (10) at the opposite end of the inlet and often passes through a mist eliminator or mist eliminator (9) to remove droplets before they are removed from the separator, as shown in the figure. Preferred mist eliminators (9) of the present invention include screens, serpentine vanes and cyclones.

The phase separator according to the present invention is usually also equipped with means for measuring and controlling the level (7) of the water phase and the level (8) of the oil phase.

The aqueous phase is discharged through the outlet (11) for water, and the oil phase through the outlet (12) for oil products. Both outlets are usually fitted with funnel quenchers to prevent funneling when the valves are opened. The vortex can potentially suck in some gas from the vapor space and draw it back into the liquid outlet.

In FIG. 6 shows a schematic diagram of another embodiment of the separation system of the invention, further comprising a flash separator for removing low boiling compounds and water from the oil phase after the second phase separator; As a rule, the flash separator operates at a temperature in the range from 80 to 150°C, for example, in the range from 100 to 130°C. The oil product is typically depressurized to near ambient before entering said flash separator, whereby the oil product is separated into: 1) a gas phase containing process gas, low-boiling oil compounds (light ends), water and, in some cases, viscosity and/or density reducers; 2) an oil phase containing dehydrated and deashed oil product. The gas from the flash separator is cooled to condense the condensable part of the gas phase, such as water, low-boiling oil fraction and/or viscosity and/or density reducers, and is further separated from the non-condensable part of the gas. The condensable portion of the gas can be further separated into an aqueous phase and an organic/light phase using gravimetric phase separation. Both the aqueous phase and the organic phase can be recycled as flushing agents in accordance with the present invention, as further shown in FIG. 8. An additional portion of the organic (light) phase may, in accordance with an embodiment of the present invention, be reblended into the oil product, as further shown in FIG. 8. Therefore, with the flash separation of the present invention, it is obtained that flushing agents can be recovered and/or the water content of the oil can be further reduced, thereby obtaining a more economical and efficient separation system.

In FIG. 7 is a schematic diagram of a preferred embodiment of the separation system of the invention further comprising a flash separator or degasser to separate gas from the converted feed mixture prior to entering the first phase separator. The flash separator or degasser of the present invention can be operated at a higher pressure than subsequent phase separators, for example, at a pressure in the range of 50-150 bar (5-15 MPa), while at least part of the process gas can be recovered at more higher pressure than downstream phase separators, resulting in an easier recovery of carbon dioxide and/or hydrogen from said gas stream. In addition, by operating said flash separator/degasser at a higher pressure than downstream phase separators, the cost of the phase

- 15 040945 outlet separators can be reduced.

In FIG. 8 shows a schematic diagram of a preferred embodiment of the separation system of the invention. The separation system includes a first phase separator for separating the product stream into a gas phase, an oil phase and an aqueous phase containing dissolved salts and liquid organic compounds of the aqueous phase. The oil phase from the first separator is further purified by mixing it with one or more flushing agents before entering the second phase separator. As shown in the figure, a preferred embodiment of the present invention may further include at least partial recycling and mixing of the separated flushing agent from the second phase separator and/or recovered light ends from the flash separator with oil. The light ends may be one or more viscosity and/or density reducers as described above. Additional additives such as additional flushing agent(s) and/or demulsifiers may be added and mixed with the oil phase as shown in the drawing.

In FIG. 9 is a schematic diagram of a preferred embodiment of a high pressure process for processing a feed stream containing carbonaceous material; the scheme includes a preferred separation system, including a recovery system for recovering liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium. The aqueous phase from the first separator and optionally the water separated in the flash stage and/or the aqueous washing agent(s) are fed to the recovery stage (6) where the process water is separated into an aqueous stream depleted in liquid organic compounds of the aqueous phase and homogeneous catalyst ( catalysts) in the form of potassium and sodium, and an aqueous stream enriched with liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium. The water stream after separation can be sent to the filtration stage before entering the extraction system (6). In addition, the pH of the aqueous stream after separation is preferably maintained in the range of 8-14, for example in the range of 9-14, preferably in the range of 10-14, for example in the range of 10-13, and in accordance with the present invention, this can be done by adding hydroxide sodium into the water stream after separation before entering the extraction unit, as shown in the figure. The outlet stream is further removed from the aqueous stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, as shown in the figure. The remaining aqueous phase, enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts, is recycled in accordance with the invention to the raw material preparation stage. The extraction unit (6) according to the present invention may include one or more methods selected from the group of evaporation, distillation, reverse osmosis, nanofiltration, ultrafiltration, pervaporation and the use of fixed beds of activated carbon.

In FIG. 10 is a schematic diagram of a preferred embodiment of the recovery system of the present invention, including an evaporation process. The process water after separation is fed to an evaporator, such as a falling film evaporator, where a fraction corresponding to the amount of water entering the high pressure system with raw materials and additives is evaporated. Typically, the ratio of concentrate to combined aqueous phases entering the recovery unit is in the range of about 0.1 to about 0.9, such as in the range of 0.2-0.8. Often the ratio of concentrate to combined aqueous phases included in the recovery unit is in the range from about 0.25 to about 0.7, for example, in the range of 0.3-0.6. In other embodiments of the present invention, the ratio of concentrate to combined aqueous phases (process water stream) entering the recovery unit is typically in the range of about 0.25 to about 0.6, such as in the range of 0.3-0.6. The process water stream after separation may be expanded in one or more flash stages before entering the evaporation stage. The process water (combined aqueous phase) may, in accordance with an aspect of the present invention, be further filtered (not shown) prior to entering the evaporator to remove, in some cases, suspended solids to reduce evaporator fouling and extend cleaning intervals. Filtration can preferably be configured to remove solid particles larger than 500 µm, for example, the filter device can be configured to remove particles larger than 250 µm; preferably, the filter device is adapted to remove particles larger than 100 microns, for example, particles larger than 50 microns. In addition, the pH of the combined aqueous phase entering the recovery according to the present invention is preferably maintained in the alkaline range, for example, in the range of 7-14, for example, in the pH range of 814, preferably the pH of the aqueous phase in the recovery unit is maintained in the range of 9-14 , for example, in the range 10-13. In an aspect of the present invention, said maintenance of pH within said range is accomplished by measuring pH and adding sodium hydroxide to the combined aqueous phase entering the recovery unit. Maintaining the pH within this range in the recovery unit has the advantage of reducing the amount of phenolics in the distillate. The evaporated fraction (distillate) can pass through the mist eliminator/foam breaker, located

- 16 040945 at the top of the evaporator, and in many applications of the present invention, the pressure of the evaporated fraction is increased somewhat, for example, by mechanical vapor recompression (as shown in the figure) or thermal vapor recompression by forcing steam into the ejector. The compression ratio can be as high as 2 in accordance with many embodiments of the present invention, for example, a compression ratio of up to 1.6 bar (0.16 MPa); preferably the compression ratio of the compressor is up to 1.3, for example up to 1.2. By increasing the pressure of the evaporated fraction, the condensation temperature of the steam is increased, so that the same steam can be used to supply the heat required in the evaporation process, making the evaporation process very energy efficient. Optionally, the recompressed vapor may be contacted with a base such as sodium hydroxide in the absorber before being returned to the evaporator on the other side of the evaporative surface. Thus, the total content of organic carbon in the distillate, for example, phenolic resins, is reduced. The alkaline absorbent solution from the absorber is preferably added to the concentrate in the evaporator and can at least partially replace the base used to maintain the pH in the evaporator. The condensed distillate may optionally undergo a coalescence step to further reduce the content of non-polar compounds before being discarded. In some cases of practice of the present invention, the condensed distillate may be further cooled and may pass through an additional final treatment stage, such as an activated carbon filter or membrane filtration, for example, a reverse osmosis stage or a bioreactor, for example, an aerobic wastewater treatment stage, before being discharged. . In this way an aqueous phase is obtained which is depleted of liquid organic compounds of the aqueous phase and of a homogeneous catalyst in the form of potassium and sodium. As shown in the figure, non-condensed vapors can be drawn off from the condensing side of the evaporator. The non-condensed vapors may, in many applications of the present invention, contain compounds having a dew point lower than that of water, such as methanol, ethanol, and acetone, as well as a non-condensable gas. According to a preferred embodiment of the invention, the non-condensed vapors may pass through an additional condenser operating at a temperature lower than the dew point of the evaporator, where the light compounds and water are further condensed as shown in the figure. The gas is separated from the additionally condensed compounds in the separator and removed. The additionally condensed compounds are preferably recycled to stage (1) of raw material preparation, preferably by mixing them with the concentrate leaving the evaporator, as shown in FIG. 10. The concentrate is preferably continuously withdrawn from the evaporator and separated into a concentrate stream leaving the evaporator and a recycle concentrate stream to the evaporator. The outlet stream according to the present invention is withdrawn from the concentrate stream as shown in the figure, and the remaining concentrate stream, enriched in liquid organic compounds of the aqueous phase and homogeneous catalyst(s) in the form of potassium and/or sodium, is recycled to stage (1) of preparing the feedstock . The outlet stream may be further processed, for example as shown in FIG. 13 and 14, or it may be incinerated or co-incinerated.

In FIG. 11 shows a schematic diagram of another embodiment of a recovery system comprising two distillation columns for separating a process water stream into a first stream enriched in aqueous phase liquid organic compounds having a boiling point lower than water, such as methanol, ethanol and acetone and water, a second a stream containing purified water for withdrawal; and a third stream containing a concentrate of aqueous phase organic compounds having a boiling point higher than water, such as phenolic compounds, water, and a homogeneous catalyst in the form of potassium and sodium. A base, such as sodium hydroxide, may be added to the process water prior to entering the first distillation column to maintain the pH of the process water within a predetermined range to control the volatility of the phenolic compounds during distillation.

In FIG. 12 shows a diagram of a preferred embodiment of a recovery unit comprising an evaporator and two distillation columns. A base, such as sodium hydroxide, may be added to the process water prior to entering the evaporator to keep the pH of the process water within a predetermined range to control the volatility of the phenolic compounds in the evaporator. Alternatively, the pH in the evaporator can be at least partially controlled by contacting the evaporated fraction with a base such as sodium hydroxide in the absorber prior to condensation and mixing said alkaline absorbent from the absorber with the process water prior to entering the evaporator and/or in the evaporator. An aqueous stream (concentrate) enriched in liquid organic compounds of the aqueous phase and a homogeneous catalyst in the form of potassium and sodium is withdrawn from the evaporator and recycled to the stage of preparation of raw materials after diverting the outlet stream to prevent undesirable accumulation of chlorides and other compounds due to this recirculation. The evaporated fraction from the evaporator is preferably condensed before entering the first distillation column in order to control the gas flow in the first distillation column. The evaporated fraction from the evaporator contains compounds lighter than water, such as methanol, ethanol and acetone, water, as well as small concentrations of heavier

- 17 040945 compounds having a boiling point higher than that of water. In the first column, the light fraction is concentrated and exits the column at the top with some water. The light fraction in accordance with the present invention can be recycled to the stage of preparation of raw materials, optionally, by mixing it with the concentrate after evaporation, after diverting the exhaust stream. The bottom product of the first distillation column is completely depleted in aqueous liquid organics and homogeneous catalysts, but is enriched in heavy compounds having a boiling point higher than that of water compared to the evaporated fraction from the evaporator, and often requires additional processing to meet environmental requirements for discharge. Accordingly, the bottoms fraction from the first distillation column is usually further processed. This treatment, according to a preferred embodiment of the present invention, can be carried out by feeding it to a second distillation column for separating into a purified aqueous product and an aqueous solution enriched in heavy organic compounds with a boiling point higher than that of water, as shown in the figure. The bottoms product may be recycled and introduced to the evaporation stage, or alternatively may be recycled to the feed preparation stage. The head product contains a purified water stream which can be discarded. In an alternative preferred embodiment of the present invention, the bottoms product from the first distillation column may be sent to a pre-discharge final treatment step, such as an activated carbon filter or a membrane process, such as a reverse osmosis or nanofiltration step, or a pervaporation step, or a bioreactor treatment. , for example, an aerobic wastewater treatment stage, instead of a second distillation column. The amount of purified water product to be discarded is equal to the amount of water entering the process along with the feedstock and other additives. Typically, this corresponds to a concentration factor of from about 1.1 to about 5, for example, a concentration factor in the range of 1.5 to 4. In addition, process water usually contains a large amount of electrolytes, such as a homogeneous catalyst, which can lead to problems deposition and pollution, as well as more frequent cleaning and maintenance; this can be difficult to control in the system of FIG. 11. The combined embodiment with evaporator and distillation is more reliable and controllable. Thus, a recovery system comprising a pH controlled evaporator producing an aqueous stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts, from which an outlet stream is withdrawn, and at least one distillation column for additional processing of the evaporated fraction from the evaporator, is a preferred embodiment of the present invention.

In FIG. 13 is a schematic diagram of a preferred exhaust stream treatment system comprising a salt separation unit comprising a first filter and two fixed beds with a chloride selective ion exchange resin. The outlet stream from the aqueous stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium (concentrate) from the evaporator is first filtered in a filter to remove suspended particles and then fed to a chloride selective ion exchange stage comprising at least two fixed beds filled with chloride-selective ion-exchange resin arranged in parallel. The valving mechanism of check valves allows the bed to be taken offline for cleaning/regeneration by backflow or backflushing with cleaning fluid while offline. Often the cleaning fluid contains deionized water. In this way, continuous operation is ensured and chloride removal can continue in the connected ion exchange bed(s) while the offline ion exchange bed(s) can be cleaned. This produces an aqueous concentrate stream with a low content of chlorides and an effluent aqueous stream enriched in chlorides. The amount of chlorides removed in accordance with the present invention is adapted to provide chloride removal corresponding to the amount of chlorides entering the process with the carbonaceous feed. Typically, the chloride removal in said ion exchange step according to the present invention is at least 50% of the chlorides in the concentrated aqueous phase entering said ion exchange step, eg the chloride removal is at least 60%. In many embodiments of the present invention, the chloride removal in said ion exchange step of the present invention is at least 70% chloride in the concentrated aqueous phase entering said ion exchange step, eg, at least 80%. The chloride-depleted stream from said chloride ion exchange stage of the present invention is preferably recycled to stage 1 of the preparation of the raw mixture, for example, by mixing it with the remaining concentrate stream from the evaporator. The chloride-rich aqueous stream is ultimately discarded after further treatment. In many embodiments according to the present invention, the amount of homogeneous catalyst(s) in the form of potassium and/or sodium, such as remaining in said chloride-depleted effluent of said chloride ion exchange stage, is at least 70 wt.% of the amount entering said stage ion exchange of chlorides, for example, at least 80 wt.%. Preferably

- 18 040945 the amount of homogeneous catalyst(s) in the form of potassium and/or sodium, such as remaining in said chloride-depleted effluent of said chloride ion exchange stage, is at least 85 wt.% of the amount entering said chloride ion exchange stage, for example, at least 90 wt.%. Therefore, less additional homogeneous catalyst needs to be added in pretreatment step 1, and an overall more efficient and economical process is provided, as will be illustrated in the following examples.

In FIG. 14 is a schematic diagram of another preferred embodiment of a recovery system including treatment of the effluent concentrate in a salt separation unit, including filtration to remove particles and at least two fixed ion exchange beds arranged in parallel, comprising a chloride selective ion exchange resin, and a valve mechanism allowing conversion to offline mode ion exchange bed for cleaning with a cleaning fluid, preferably deionized water. As shown in the figure, the processing of the effluent concentrate further includes a second effluent taken from the chloride-depleted concentrate stream downstream of the salt separation unit. The second exhaust stream is diverted to prevent sodium accumulation in the system.

Example 1

Obtaining water concentrate.

Fuelwood (a mixture of mainly Scandinavian spruce, pine, birch, including bark) with a moisture content of 36.6 wt.% and a chloride content of 74 mg/kg was ground in a hammer mill to obtain a maximum particle size of 1 mm and mixed in a mixer with high shear rate with a recycled aqueous concentrate comprising liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, with recycled liquid organic compounds of the oil phase, an additional catalyst in the form of potassium carbonate and sodium hydroxide to obtain a raw mixture containing the following.________________________________

Component % wt. Chopped fuel wood (by dry weight) 23.0 Recycled Liquid Organic Oil Phase Compounds 23.0 Water 48.0 Liquid organic compounds of the aqueous phase 4.44 Potassium 0.91 Sodium 0.65 Chloride 0.006

The feed mixture was converted in a continuous plant by pressurizing it to 334 bar (3.34 MPa) and heating it to a temperature of 402°C, and keeping the feed mixture under conversion conditions for approximately 12 minutes before cooling it to 97°C, filtering through a 500 µm stainless steel filter, expanding and further cooling the converted feed mixture to ambient pressure and a temperature of 63 °C through a pressure reduction system containing a number of tubular elements and an additional cooler, and degassing the product in a degasser. The liquid phases were manually separated into an oil phase containing liquid organic compounds of the oil phase and an aqueous phase containing liquid organic compounds of the aqueous phase and homogeneous catalysts in the aqueous phase containing liquid organic compounds of the aqueous phase.

The aqueous phase was subjected to a recovery process as shown in FIG. 10, where the concentration factor was approximately 2.2. The extraction system operated in four different configurations, as shown in Table 1 below. 1. The concentrate shown is a combined concentrate for all four configurations.

- 19 040945

Table 1

Operating mode/Parameter Process water Water concentrate Output stream water Outgoing water flow Outgoing water flow Outgoing water flow Evaporator X X X X Absorber X X X coalescer X X AC filter X X pH 8.8 10.2 8.7 9.2 8.2 7.8 Na, g/l 16 34 <0.0005 <0.0005 <0.0005 <0.0005 K, g/l 22 48 <0.0005 <0.0005 <0.0005 <0.0005 C1, mg/l 155 340 not def. not def. not def. not def. pH 8.8 10.2 8.7 9.2 8.2 7.8 TOC, g/l 56 134 1.9 2.3 0.0019 0.0012 Methanol, g/l 6.2 0.75 1.7 2 <0.2 0.0058 Ethanol, g/l 3.6 0.2 1.2 0.72 <0.2 0.0011 Acetone, mg/l 370 6.7 39 20 0.78 0.79 Phenols, mg/l >20 >8.5 >25 <0.42 <0.0001 <0.0001

As can be seen from Table. 2, almost complete extraction of potassium and sodium from the condensate in the evaporator was obtained. In addition, it can be seen that most of the TOC remains in the concentrate (98.5%). However, most of the lighter components, such as methanol, ethanol, and acetone, end up in the vaporized fraction, as shown in the table, and are only recovered in configurations containing an activated carbon filter. It was not possible to measure the concentration in both process water and concentrate. However, as can be seen from the table, the amount of phenols is significantly reduced by the alkaline absorber and is reduced below the limit of detection for configurations containing an activated carbon filter.

Example 2

Removal of chlorides from water concentrate.

The aqueous phase concentrate in Example 1 was sent to the effluent treatment system as shown in FIG. 13, with the results presented in table. 2.

table 2

Process water Concentrate cleansing water Chlorine-depleted stream Chloride-rich stream Flow, % 100 120 100 120 Na, g/l 72.0 0 67.6 3.8 K, g/l 48.0 0 44.7 2.8 С1, g/l 340 0 67 227 pH 10.2 not def. not def. not def.

As can be seen from the table, the removal of chlorides is approximately 80%, and the extraction of potassium and sodium is 93.2 and 93.3%, respectively. The distribution of TOC in different streams was not measured. It should be noted that while the effluent treatment system can control the chloride concentration, other compounds such as sodium can accumulate if the process is performed as shown in FIG. 13 without further diverting the exhaust stream as shown in FIG. 14 and further illustrated below in Example 3.

Example 3

The need for an exhaust stream.

In table. 3 compares the effluent requirements to prevent build-up of impurity elements and the degree of addition of potassium and sodium based on the data in Examples 1 and 2 and a chloride threshold of 400 mg/L in the effluent diversion concentrate without and with chloride selective ion exchange treatment. in accordance with FIG. 14.

Claims (35)

Table 3 No exhaust flow treatment With exhaust flow treatment Required amount of flushing CONC, water, % 15.0 5.9 conc. chlorides in conc. water flow, ppm 400 377 Amount of additional catalyst, % 15.0 5.9 Amount of additional base, % 15.0 5.9 As can be seen from Table. 3, the effluent requirements, and therefore also the potassium and base addition requirements, are significantly reduced by carrying out the effluent treatment. This represents the operating costs for the main streams added. In addition, as the demand for the exhaust stream decreases, the amount of aqueous phase liquid organic compounds recycled to the feed preparation stage also increases, and the overall oil yield increases accordingly. CLAIM 1. A method for separating and purifying products from a high-pressure processing system designed to process a raw mixture containing carbon-containing material (materials) at a pressure of 150 bar (15 MPa) to 400 bar (40 MPa) and a temperature of 300 to 430 ° C in the presence of homogeneous catalysts in the form of potassium and/or sodium in a concentration of at least 0.5 wt.% and liquid organic compounds in a concentration of 5 to 40 wt.% for a predetermined time, resulting in the formation of a converted raw mixture, where the converted the raw mixture is cooled to a temperature in the range of 50 to 250°C and its pressure is reduced to a pressure in the range of 1-150 bar (0.1-15 MPa), and where the converted raw mixture is separated into a gas phase containing carbon dioxide, hydrogen and methane, an oil phase containing liquid organic compounds of the oil phase, and an aqueous phase containing liquid organic compounds of the aqueous phase, dissolved salts, where liquid organic compounds the aqueous phase and dissolved homogeneous catalysts in the form of potassium and/or sodium are at least partially removed from said aqueous phase, resulting in a first aqueous phase stream enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, and a second aqueous phase stream depleted of aqueous phase liquid organic compounds and homogeneous catalysts in the form of potassium and sodium, and wherein the first aqueous phase stream is at least partially recycled to said feed mixture to produce at least a portion of said liquid organic compounds and homogeneous catalysts in the feed mixture, and where, in addition, before recycling said first recirculation stream to the feed mixture, an outlet stream is diverted from said aqueous phase enriched in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and sodium, and the outlet stream is additionally treated for one and more ion exchange stages, and said one or more ion exchange stages comprise one or more ion exchange resins contained in one or more fixed beds in a parallel arrangement with check valves before and after each layer, whereby at least one ion exchange bed is connected and one or more ion exchange beds can be transferred at least offline for cleaning, thereby allowing continuous operation. 2. The method according to claim 1, in which the amount of discharged exhaust stream is chosen so as to obtain a concentration of chloride in the raw mixture of less than 600 hours/million by weight, for example less than 400 hours/million by weight; preferably less than 200 ppm by weight, for example less than 100 ppm by weight. 3. The method according to any one of claims 1, 2, in which the concentration of chloride in the first stream of the aqueous phase is less than 250 hours/million by weight, for example less than 200 hours/million by weight; preferably the chloride concentration in the first aqueous phase is less than 150 ppm by weight, for example less than 100 ppm by weight. 4. Process according to any one of claims 1 to 3, wherein the weight ratio of said bleed off stream to the first aqueous phase stream is in the range of 0.01-0.4, for example in the range of 0.02-0.25, preferably the weight ratio said discharged outlet stream to the first aqueous phase stream is in the range of 0.03-0.2, for example in the range of 0.04-0.15. 5. The method of claim 1 wherein the ion exchange resin comprises a chloride selective resin. 6. The method of claim 5, wherein the ion exchange bed(s) is further provided with a valve mechanism to regenerate/clean said ion exchange resins by backflowing and/or backwashing the cleaning fluid while offline. - 21 040945 7. The method of claim 6, wherein the cleaning fluid contains demineralized water. 8. A process according to any one of claims 5 to 7 wherein the effluent stream is filtered to remove suspended particles prior to entering said ion exchange step(s). 9. The method according to any one of claims 5, 6, in which the pH value of the outlet stream at the inlet to the stage of ion exchange is maintained in the range from 8 to 14; preferably in the range of 9-14 and even more preferably in the range of 10-14. 10. A process according to any one of the preceding claims, wherein the organic liquids in said feed mixture comprise recycled oil phase organic liquids. 11. The method according to any one of the preceding claims, wherein the pressure of the converted raw mixture is reduced to a value in the range of 10-50 bar (1-5 MPa) before said separation. 12. The method according to any one of the preceding claims, wherein the aqueous phase entering the recovery system comprises aqueous phase liquid organic compounds having a boiling point less than water and aqueous phase liquid organic compounds having a boiling point greater than water . 13. The method according to claim 12, in which the liquid organic compounds of the aqueous phase contain one or more components selected from one or more of the following groups: a) ketones such as acetone, propanones, butanones, pentanones, pentenones, cyclopentanones such as 2,5-dimethylcyclopentanone, cyclopentenones, hexanones and cyclohexanones, for example 3methylhexanone, quinones; b) alcohols and polyalcohols such as methanol, ethanol, propanols, butanols, pentanols, hexanols, heptanols, octanols such as 2-butyl-1-octanol, hydroquinones, benzene diols; c) phenols, alkylated phenols, polyphenols, monomeric and oligomeric phenols, creosol, thymol, alkoxyphenols, para-coumaryl alcohol, coniferyl alcohol, sinapyl alcohol, flavenols, catechins; d) carboxylic acids such as formic acid, acetic acid, and phenolic acids such as ferric acid, benzoic acids, coumaric acid, cinnamic acid, abietic acid, oleic acid, linoleic acid, palmitic acid, stearic acid; e) furans such as tetrahydrofuran (THF); f) alkanes, alkenes, toluene, cumene. 14. The method according to claim 13, wherein the concentration of individual liquid organic compounds of the aqueous phase in the aqueous phase entering the recovery system is less than 2.0 wt.%, for example, less than 1.0 wt.%. 15. A method according to any one of the preceding claims, wherein the aqueous phase organic liquid compounds comprise emulsified oil phase droplets. 16. The method according to any one of the preceding claims, wherein the recovery of aqueous phase liquid organic compounds and homogeneous catalysts in the form of potassium and/or sodium from the aqueous phase comprises one or more of evaporation, distillation/fractionation, reverse osmosis, nanofiltration, ultrafiltration and pervaporation. 17. The method according to claim 16, in which the recovery of liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and / or sodium from the aqueous phase includes one or more stages of evaporation and / or distillation, resulting in the formation of a first aqueous phase enriched in liquid aqueous phase organic compounds and homogeneous catalysts in the form of potassium and/or sodium (concentrate), and a second aqueous phase stream depleted in liquid organic compounds of the aqueous phase and homogeneous catalysts in the form of potassium and/or sodium (distillate), where the amount of the resulting second aqueous phase is chosen to match the amount of water entering the high pressure treatment system, eg contained in one or more carbonaceous feedstocks. 18. The method of claim 17 wherein the aqueous phase entering said recovery system is filtered to remove suspended solids prior to entering said one or more evaporation and/or distillation stages. 19. The method of claim 17 wherein the recovery system comprises one or more flashing stages. 20. The method according to claim 17, wherein the pH of the aqueous phase in the recovery system is maintained in the alkaline range, for example in the pH range of 7-14, preferably the pH is maintained in the range of 9-14, for example in the range of 10-13. 21. The method of claim 20, wherein said alkaline maintenance comprises measuring and adjusting the pH by adding sodium hydroxide to the aqueous phase. 22. The method of claim 19, wherein the recovery system comprises at least one evaporator, such as a falling film evaporator. 23. The method according to claim 17, wherein the evaporated vapor of said evaporation stage is condensed in at least two condensation stages having a decreasing temperature. 24. The method according to claim 23, in which the evaporated vapor passes through the eliminator and / or coalescer - 22 040945 before said condensation stage. 25. The method of claim 23 wherein the evaporated vapor is brought into contact with an absorbent in the absorber prior to said condensation steps. 26. The method of claim 25 wherein the absorbent comprises an alkaline absorbent. 27. The method of claim 26 wherein the absorbent comprises sodium hydroxide. 28. The method of claim 16, wherein the recovery system comprises one or more distillation columns, each equipped with a desorption and distillation section. 29. The method of claim 1, wherein the separation of the converted feed mixture is carried out in a first phase separator operating at a temperature of 50 to 250°C, wherein the oil phase from the first phase separator is further purified by mixing it with one or more flushing agents and wherein the oil phase is separated from one or more flushing agents in a second phase separator operating at a temperature of from 50°C to about 250°C. 30. The method according to claim 29, wherein the first and/or second phase separator is operated at a pressure in the range of 10-100 bar (1-10 MPa). 31. The method of claim 29, wherein the separation of the oil and flushing agent(s) also includes one or more flash distillation steps of the oil phase and/or separated flushing agents downstream of the second phase separator. 32. The method of claim 31, wherein the separation also includes a second washing step in which the separated oil phase from the second phase separator is mixed with one or more additional washing agents and separated from the one or more additional washing agents in a third separator located between the second phase separator and the flash distillation stage. 33. The method of claim 29, wherein one or more water-containing streams are withdrawn from the second phase separator and/or the third separator and/or the flash distillation stage and mixed with the separated water from the first separator prior to entering the recovery system and /or inside the extraction system. 34. The method of claim 1, wherein the one or more carbonaceous materials are selected from the following: biomass, such as woody biomass and waste, such as wood chips, sawdust, timber harvested from thinning, road construction, bark, branches, garden and park waste and weeds; energy crops such as coppice, willow, miscanthus and reed arundo; agricultural residues and by-products such as grasses, straw, stalks, husks, corn cobs and hulls, eg wheat, rye, corn, rice, sunflower; empty fruit bundles from palm oil production, palm oil production effluents (ROME), residues from sugar production such as bagasse, vinasse, molasses, greenhouse waste; energy crops such as miscanthus, millet, sorghum, jatropha; aquatic biomass such as macroalgae, microalgae, cyanobacteria; livestock bedding and compost, such as the fibrous fraction of feed production; municipal and industrial waste streams, eg black liquor, paper sludge, off-grade pulp and paper fibers; residues and by-products from food production such as juice or wine production, vegetable oil production, sorted municipal solid waste, sorted household waste, restaurant waste, meat processing waste, sewage sludge, plastics, and combinations thereof. 35. The method of claim 1 wherein the aqueous phase contains optionally suspended particles. -