DESCRIPTION

BACKGROUND

Ligno-cellulosic biomass may be used to produce a platform of valuable chemicals and biofuels derived from monomeric sugars, such as glucose and xylose. Recovery of monomeric sugars from the insoluble carbohydrates of the ligno-cellulosic biomass (cellulose and hemicellulose) is a multistep process usually comprising a pre- treatment of the ligno-cellulosic biomass to increase the accessibility to the polysaccharides, which must then be cleaved to monosaccharides in a process known as hydrolysis.

Different physical, thermal, biological and chemical pretreatments have been developed so far. Most effective pretreatments involve hydrothermal processes often in the presence of aggressive chemical agents, such as sulfuric acid, which solubilizes a portion of the hemicellulose.

Hydrolysis of pre-treated ligno-cellulosic biomass can be conducted either as mineral acid catalyzed or as enzymatically catalyzed process.

Enzymatic hydrolysis is a selective process involving enzymes as catalyst, wherein very few by-products are formed. Enzymes are sensitive to inhibitory products which are formed during the pre- treatment of the ligno-cellulosic biomass.

In acid hydrolysis, both cellulose and hemicellulose chains are cut into monosaccharides. Acid hydrolysis can be catalyzed with either dilute or concentrated acid, usually with sulfuric acid as catalyst. The negative aspects of acid-catalyzed hydrolysis are high acid consumption and corrosion problems in the concentrated acid hydrolysis, and formation of degradation products mainly in the dilute acid hydrolysis. Moreover, the added acid increases significantly the amount of anionic species, which have to be removed before subsequent conversion steps of the monomeric sugars increasing the complexity and the cost of purification step. Processing the ligno-cellulosic biomass to remove contaminants is known.

WO201 1028554 discloses a method for pretreatment of particulate cellulosic biomass feedstock, wherein the method comprises the step of removing from the biomass feedstock a fine particulate fraction, which is solid, thereby forming a cleaned particulate biomass feedstock.

The cleaned particulate biomass feedstock has an acid neutralization capacity of less than 0.01 , or alternatively that is no more than about 90% of the acid neutralization capacity of the biomass feedstock, as determined in accordance with protocols defined in the application. The application discloses copious embodiments to further processing the cleaned particulate biomass feedstock by means of acidic impregnation, steam treatment, and enzymatic hydrolysis.

WO20101 13129A2 discloses another process for the treatment of ligno- cellulosic biomass, wherein a ligno-cellulosic biomass feedstock is soaked in vapor or liquid water or mixture thereof in the temperature range of 100 to 210 °C for 1 minute to 24 hours to create a soaked biomass and a first liquid. At least a portion of the first liquid is separated from the soaked biomass to create a first liquid stream and a first solid stream, wherein the first solid stream comprises the soaked biomass. The first solid stream is steam exploded to create a steam exploded stream comprising solids and a second liquid. The first liquid stream may then be combined with the steam exploded stream and subjected to enzymatic hydrolysis. In an embodiment, the soaking step is preceded by a low temperature soaking step, wherein the ligno- cellulosic biomass is soaked in liquid comprised of water at a temperature in the range of 25 to 100°C for 1 minute to 24 hours and the low temperature soaking step is followed by a separation step to separate a least a portion of the liquid from the low temperature soak.

WO2013098789A1 discloses an improved method of conducting a pre- soaking step involving pre-soaking the ligno- cellulosic biomass in a liquid (water) at a temperature in the range of between 100°C to 150°C prior to soaking at higher temperatures thereby producing a soaking liquid and a solid stream comprising the biomass.

The soaked liquid may then be filtered by nano-filtration to separate acetic acid. It disclosed that higher temperatures work better for removing the contaminants, particularly the contaminants which inhibit nano-filtration of the liquid from the soaking step, even if temperatures greater than 100°C require pressure vessels, special heating kits, insulation and increase the capital and operating cost of the process. When nano-filtration is used, the pre-soak temperature can be in the range of 10°C to 150°C.

The three applications cited above disclose a purification or conditioning step or steps of the pre-treated feedstock before the enzymatic hydrolysis of the pre treated feedstock to remove inhibitors. The application do not disclose any conditioning of the hydrolyzed streams by means of chromatographic separation to produce a purified liquid sugar stream.

In WO2013026849 it is disclosed a process for the hydrolysis of oligosaccharides present in a liquid biomass feed stream derived from pre-treated ligno-cellulosic biomass. The process comprises creating an acidic stream from the liquid biomass feed stream by increasing the number of H ⁺ ions to the liquid biomass feed stream in an amount sufficient so that the pH of the acidic stream is at least 0.5 pH units less than the pH of the liquid biomass feed stream prior to the addition of the H ⁺ ions wherein less than 80% of the total amount of H ⁺ ions added to the feed stream are derived from an acid or acids, and then hydrolyzing the acidic stream by increasing the temperature of the acidic stream to a hydrolysis temperature greater than 80 °C.

In a preferred embodiment, at least a portion of the H ⁺ ions come from decationization using an ion exchange agent and at least a portion of the acidic stream is separated from the ion exchange agent before hydro lyzing the separated portion of the acidic stream.

The application does not recognize any specific advantage enabled by the disclosed hydrolysis processes for recovering a purified monomeric sugar stream from the hydrolyzed acidic stream and therefore it does not disclose any purification configuration of the hydrolyzed acidic stream.

As a result of the severity of the pre-treatment and the hydrolysis processes, it is obtained a ligno-cellulosic biomass hydrolyzate which comprises the desirable monomeric sugars, but also unwanted soluble compounds derived both from the ligno-cellulosic biomass and from the process. In particular, a large amount of impurities, including dissociated salts, sugar degradation products, organic acids, soluble phenolic compounds, and other compounds are present in the sugar stream after the hydrolysis. These soluble compounds result from degradation or solubilization of the feedstock or from the acids and alkali added in the process. As many of these soluble compounds will limit the following conversion of the monosaccharides to the desired chemicals and biofuels, the hydrolyzate must be refined to produce a purified monomeric sugars liquid stream.

The recovery and purification of monomeric sugar streams from the ligno-cellulosic biomass hydrolyzate has been the subject of a significant amount of research. The processes studied include lime addition and chromatographic separation, which may be conducted by ion exchange and by ion exclusion.

In lime addition, lime (calcium hydroxide), which is insoluble, is added to the sugar stream to precipitate impurities. The limed sugar solution has an alkaline pH and is neutralized with acid, typically phosphoric acid, sulfurous acid, carbonic acid, or a mixture thereof. Optionally, the lime cake is separated from the sugar by filtration. A second option is to filter the lime cake at alkaline pH and carry out a second filtration to remove material that precipitates during the acidification steps. Lime treatment decreases the toxicity of the sugar stream to yeast and other microbes. However, any handling of the lime cake is difficult and costly. In addition, the introduction of calcium into the stream increases the likelihood that calcium scale will deposit on evaporators, distillation columns, and other process equipment. The clean-up and avoidance of scale increases the cost of sugar processing. Furthermore, the introduction of lime makes the recovery of salt and acetic acid more difficult.

It is known in the art that chromatography and adsorption, or a combination thereof, may be used to produce food grade, high value added monomeric sugar streams. An extensive review of the purification and fractionation processes of ligno-cellulosic hydrolyzates may be found in Jari Heinonen, Tuomo Sainio, "Chromatographic Fractionation of Lignocellulosic Hydrolysates", Advances in Chemical Engineering, Volume 42, pag.261.

In the case of bio-fuel or low added-value chemicals, the recovery of monomeric sugar streams must be accomplished in a cost-effective way.

In order to improve the economics of the acid-catalyzed hydrolysis processes, the hydrolysis acid should be recovered and recycled. In the chromatographic fractionation of concentrated acid ligno-cellulosic hydrolyzates, strong cation-exchange resins in acid form are used as an adsorbent. The hydrolyzates formed in the acid hydrolysis process are fractionated into hydrolysis acid fraction, monosaccharide fraction, and by-products fractions. The hydrolysis acid fraction is led through a concentration step back to the hydrolysis. This recycling lowers the chemical consumption (costs) of the process considerably because the need for fresh acid is reduced. One of the problem related to the chromatographic separation of hydrolyzates from acid hydrolysis process is the presence in the hydrolyzate of anions from the added mineral acid. Moreover other ionic species from the biomass are present in the hydrolyzate, such as cations and anions originated from salts contained in the biomass and dissolved during the pretreatment and hydrolysis steps.

As evidenced for instance in figure 5.10 of the review of Jari Heinonen at al., reproducing the chromatographic diagram of a concentrated acid spruce hydrolyzate, a huge peak comprising charged species is overlapped to the monomeric sugars peak. Thereby, the separation of a monomeric sugars stream from an acid hydrolyzate with high purity and yield may not be accomplished by means of a unique chromatographic step, while the use of many chromatographic steps will increases the cost of the separation process. In ion exchange, the sugar stream is flowed through columns packed with ion exchange resins. The resins are in a cation exchange or anion exchange form, or a combination of the two. In principle, cation- exchange resins remove cations such as sodium or potassium, while anion-exchange resins remove anions such as sulfate and acetate. For example, ion exchange has been investigated by Nilvebrant et al. (App. Biochem. Biotech., 2001 , 91-93:35-49) in which a spruce hydrolyzate was treated to remove inhibitors, such as phenolic compounds, furan aldehydes and aliphatic acids. The separation was carried out using an anion exchanger, a cation exchanger and a resin without charged groups. The investigators found that a treatment at pH 10.0 using an anionic exchanger removed phenolic compounds since at this pH most of the phenolic groups were ionized.

In practice, several factors limit the effectiveness of ion exchange treatment to remove inhibitors. First, the multi-component nature of the streams results in an inefficient removal of some species, at any single set of conditions. Second, the high ionic load demands very frequent and expensive regeneration of the resin. Finally, not all of the inhibitors are ionic, and ion exchange is ineffective in removing nonionic compounds from sugar.

Ion exclusion uses ion exchange resins, but rather than bind target ions in solution, the charge on the resin matches that of the target ions in the solution, thereby excluding them from the resin.

The excluded compounds then elute from the column readily, while uncharged compounds absorb into the resin and elute from the column more slowly. For example, a concentrated solution of sulfuric acid and glucose has hydrogen as the primary cation. A cation-exchange resin in the hydrogen form will exclude the acid, causing it to elute quickly. The glucose, which is uncharged, is not excluded from the resin and absorbs into the resin void, thereby eluting from the column more slowly than the acid.

Ion exclusion has been used for example by Wooley et al., (Ind. Eng. Chem. Res., 1998, 37:3699-3709), which teaches the removal of acetic acid and sulfuric acid from biomass sugars by pumping a product stream over a bed of cation exchange resin in the hydrogen form.

The positive charge on the resin repels the hydrogen ion in the sulfuric acid, thereby causing the sulfuric acid to elute from the column very quickly. The uncharged sugar molecules are absorbed into the void space of the resin and elute from the column more slowly than the sulfuric acid. Fully associated acetic acid (non-ionic) is a smaller molecule than sugar or sulfuric acid and so elutes from the column more slowly than sulfuric acid or sugar. Also described is a Simulated Moving Bed (SMB) system for producing a glucose stream free of sulfuric acid and acetic acid. The ion exclusion was carried out at a pH of between about 1-2 and, at such low pH values, significant degradation of xylose is likely.

US5560827 and US5628907 disclose a process for separating an ionic component (acid) from a non-ionic component (sugar) using an SMB arrangement, including a plurality of ion exclusion columns arranged in 4 zones. The separations are run at a low pH using a cationic (or cation- exchange) resin in the hydrogen form. The methods incorporate various arrangements to minimize the dispersion and channeling effects. The sugar/ acid solution is loaded onto the column and the acid elutes first while sugar is eluted later using water.

US5407580 discloses a process for separating an ionic component (acid) from a non-ionic component (sugar) using a preparative-scale ion exclusion system. The system includes a floating head distribution plate to prevent evolution of a dilution layer caused by the shrinkage of the resin bed. The columns can be operated over a range of process conditions to produce separate and distinct elution profiles for the acid and sugar. Acceptable conditions for carrying out the process are at a sulphuric acid concentration of 1.0 to 20.0%, a feed volume of 1.0 to 5.0, a flux rate of 0.1 to 2.0 and using a divinylbenzene resin with a percent crosslinking of between 1.0 and 15.

US5968362 discloses a method of separating sugars and acid by ion exclusion chromatography using an anion exchange resin. The sugars elute through the column, and may contain residual acid and heavy metals. The heavy metals, can be removed and the acid neutralized using a lime treatment. The acid adsorbs to the resin and is retained; it is eluted from the resin with water.

US6663780 discloses a method in which product fractions, such as sucrose, betaine and xylose, are separated from molasses that are obtained from a variety of sources, including beet and cane molasses, as well as hydro lyzates produced from biomass. The process involves treating the molasses with sodium carbonate (pH 9) to precipitate calcium followed by removing the resulting precipitate. The filtrate is then subjected to a simulated moving bed (SMB) process which is carried out using at least two SMB systems packed with a strongly acid cation exchange resin. Sucrose is recovered in a first system and betaine is recovered in a second system. The sucrose obtained from the first system may be crystallized and the crystallization run-off applied to the second system. Also, described is a process for recovering xylose from sulphite cooking liquor using two systems. Prior to fractionation in the first system, the sulphite cooking liquor, having a pH of 3.5, is filtered and diluted to a concentration of 47% (w/w). The xylose fractions obtained from the first system are crystallized and, after adjustment to pH 3.6 with MgO, the run-off is fed to the second system. In the second system, a sequential SMB is used to separate xylose from the crystallization run-off. A disadvantage of the separation technique disclosed in US6663780 is that the inclusion of two SMB systems is costly and adds to the complexity of the process. Moreover, the initial sucrose purification by crystallization is an expensive technique.

WO9517517 discloses a method of processing municipal solid waste to recover reusable materials and to make ethanol. Cellulosic material is shredded and pre-treated with acid and lime to remove heavy metals, then treated with concentrated acid (sulfuric) to produce sugars. The sugars and the acid are separated on a strong acidic cation ion exchange resin. US4101338 discloses a method of separating salts and sucrose present in blackstrap molasses obtained from sugar cane by ion exclusion chromatography. Prior to ion exclusion chromatography, the molasses are treated by removing organic non-sugar impurities and colour. Various methods are suggested for removing these impurities, including a preferred method utilizing precipitation with iron salts, such as ferric chloride or ferric sulfate, to form floes. The insoluble floes are then removed from the molasses stream and the soluble iron salts are removed by the addition of lime and phosphoric acid or inorganic phosphate salts, thereby raising the pH to above 7.0. The molasses stream is then applied to the ion exchange column to produce fractions containing sucrose and separated salts. A disadvantage of this process is that, upon addition of ferric ions, the molasses has a pH that is in the range of 2.0 to 3.0. At such a low pH, degradation of xylose could occur.

US4631 129 teaches a method of purifying sugar from a sulfite pulping spent liquor stream. The process involves two steps, in which, during the first step, the pH of the spent sulfite liquor is adjusted to below 3.5 and the stream is passed through a strongly acidic ion exclusion resin to recover two lignosulfonate-rich raffinate fractions and a product stream containing the sugar and consisting of 7.8%-55% lignosulfonate. In the second step, the product stream is adjusted to pH 5.5-6.5. The product stream is then filtered, and applied to a second ion exclusion column to further purify the sugar by separating it from the large amount of lignosulfonates in this stream. A problem with this process is that the use of two ion exclusion systems is costly and adds to the complexity of the process. Moreover, the patent does not quantify or address the separation of compounds present during the processing of biomass such as inorganic salts, including sulfate salts, and acetic acid and other organic acids.

US8003352 discloses a process for obtaining a product sugar stream from cellulosic biomass, which comprises the main steps of pretreating the cellulosic biomass at a pH of about 0.4 to about 2.0 by adding one or more than one acid to the cellulosic biomass to produce a pretreated cellulosic biomass; adding one or more than one base to the pretreated cellulosic biomass to adjust the pretreated cellulosic biomass to a pH of about 4.0 to about 6.0, thereby producing a neutralized cellulosic biomass comprising inorganic salt and acetate salt; hydrolyzing the neutralized cellulosic biomass with cellulase enzymes to produce a crude sugar stream; separating insoluble residue from the crude sugar stream to produce a clarified sugar stream; treating the clarified sugar stream using ion exclusion chromatography with a cation exchange resin at about pH 5.0 to about 10.0 to produce one or more than one raffinate stream comprising the inorganic salt and acetate salt and a product sugar stream comprising sugar. The disclosed process adds thereby a large amount of anionic species which must be removed by chromatography, and then use of enzyme, both the steps increasing considerably the cost of the process.

Although the production of a ligno-cellulosic hydrolyzate and the recovery of a purified stream of monomeric sugars from the ligno- cellulosic hydrolyzate has been extensively studied, the processes developed so far are characterized by high costs and complexity.

Thereby there is still the need to develop a cost-effective process to produce a stream of monomeric sugars from a ligno-cellulosic biomass, said stream having a high purity and said process being characterized by a high sugar recovery yield.

SUMMARY

It is disclosed a process for producing a liquid sugar stream comprising monomeric sugars from a ligno-cellulosic biomass stream, comprising a ligno-cellulosic component and non-ligno-cellulosic water soluble compounds including water soluble salts. The process comprises the steps of:

Subjecting the ligno-cellulosic biomass stream to a first hydrothermal treatment corresponding to a first severity factory which is derived by treating the ligno-cellulosic biomass stream at at least a first hydrothermal temperature for at least a first hydrothermal time, to produce a presoaking liquid stream and a presoaked ligno-cellulosic biomass stream, wherein the presoaking liquid stream comprises water and dissolved species derived from the non-ligno-cellulosic water soluble compounds; treating the presoaked ligno-cellulosic biomass to produce a liquid biomass feed stream and a treated ligno-cellulosic biomass stream, wherein the liquid biomass feed stream comprises water, oligomeric sugars derived from the ligno-cellulosic component, and additional dissolved species derived from the non-ligno-cellulosic water soluble compounds, comprising disassociated cations and anions of the water soluble salts; hydrolyzing at least a portion of the oligomeric sugars of the liquid biomass feed stream to monomeric sugars to create a liquid sugar stream, wherein the hydrolysis is conducted at a pH lower than 5.5; raising the pH of the liquid sugar stream to a pH in the range of 5.5 to 8, by adding a sufficient amount of a reference base comprised of a reference cation and a reference anion to form a stoichiometric amount of a reference salt comprised of the reference cation and the disassociated anion, wherein the ratio of the total equivalents of cations different from the reference cations to the total equivalents of cations in the liquid sugar stream is less than 0.20; separating the liquid sugar stream into at least the purified liquid sugar stream and a first residual stream by means of a cationic resin, wherein the cation of the cationic resin is the same of the reference cation, wherein the purified liquid sugar stream is characterized by a purity ratio of the total amount of monomeric sugars to the total amount of all the compounds on a dry basis, and the purity ratio of the purified liquid sugar stream is greater than 60%.

It is also disclosed that the hydrolysis may comprise the steps of adding to the liquid biomass feed stream an amount of H ⁺ ions sufficient to create an acidic liquid biomass feed stream having a pH which is less than a value selected from the group consisting of 3, 2.5, 2, and 1.5; and increasing the temperature of the acidic liquid biomass feed stream in the range of 80 °C to 200 °C.

It is further disclosed that the majority of the H ⁺ ions may be derived from a decationization of the liquid biomass feed stream using an ion exchange agent.

It is also disclosed that the percent of the total amount of H ⁺ ions derived from an acid or acids added to the liquid biomass feed stream may be less than a value selected from the group consisting of 50%, 40%, 30%, 20%, 10%, and 1%.

It is further disclosed that the ratio of the total equivalents of cations different from the reference cations to the total equivalents of cations in the liquid sugar stream after the addition of the reference base may be less than a value selected from the group consisting of 0.15, 0.10, 0.05 and 0.03.

Alternatively, it is also disclosed that the hydrolysis may comprise contacting the liquid biomass feed stream with an enzyme or enzyme cocktail, and that the enzyme or enzyme cocktail may comprise an hemicellullase.

It is further disclosed that the purity ratio of the purified liquid sugar stream may be greater than a value selected from the group consisting of 70%, 80%, 85% and 90%.

It is also disclosed that the recovery ratio of the separation process may be greater than a value selected from the group consisting of 70%, 80%, 85%, 90%, and 95%.

It is further disclosed that the cationic resin may have an ion-exclusion action, and preferably also a size-exclusion action.

It is also disclosed that the reference cation is a monovalent cation, and that preferably the monovalent cation is Na ⁺ .

It is further disclosed that the reference base is preferably sodium hydroxide. It is also disclosed that the ratio of the total equivalents of non- monovalent cations to the total equivalents of cations in the liquid sugar stream after the addition of the reference base may be less than a value selected from the group consisting of 0.05, 0.03, 0.02 and 0.01.

It is further disclosed that the separation of the purified liquid sugar stream may occur in a continuous mode by means of at least a technique selected from the group consisting of Simulated Moving Bed, Improved Simulated Moving Bed, Sequential Simulated Moving Bed.

It is also disclosed that the purified liquid sugar stream may further be subjected to a refining step to remove at least a portion of the soluble compounds which are not monomeric sugars to produce a refined liquid sugar stream, wherein the refined liquid sugar stream has a purity ratio of the refined liquid sugar stream which is greater than a value selected from the group consisting of 90%, 95%, 97%, 98%, 99% and 99.9%.

It is further disclosed that the refining step is further characterized by a recovery ratio of the refining step which is the ratio of the total amount of monomeric sugars in the refined liquid stream to the total amount of monomeric sugars in the purified liquid sugars stream, and the recovery ratio of the refining step may be greater than a value selected from the group consisting of 95%, 97%, 98% and 99%.

It is also disclosed that the refining step may comprise contacting the purified liquid sugar stream with at least a refining material selected from the group consisting of a cationic resin, an anionic resin and a carbonaceous material, and that the carbonaceous material preferably comprises an activated carbon.

It is further disclosed that the refining step comprises first contacting the purified liquid sugar stream with at least a first refining material selected from the group consisting of a cationic resin, an anionic resin, or a combination thereof, and then further contacting with a second refining material comprising an activated carbon.

It is also disclosed that the monomeric sugars may comprise C5 monomeric sugars, and the percent amount of C5monomeric sugars on the total amount of monomeric sugars in the refined liquid sugar stream may be greater than a value selected from the group consisting of 50%, 60%, 70%, 80%, and 90%.

It is further disclosed that the first hydrothermal treatment may be conducted at at least a temperature in a range selected from the group consisting of 30°C to 100°C, 40°C to 99°C, 40°C to 90°C, and 50°C to 85°C. It is also disclosed that the first hydrothermal treatment may be conducted for at least a time in a range selected from the group consisting of 10 seconds to 300 minutes, 1 minute to 20 minutes, 2 minutes to 20 minutes, 2 minutes to 15 minutes, and 3 to 10 minutes. It is further disclosed that the treatment of the presoaked biomass may comprise subjecting the presoaked ligno-cellulosic biomass stream to a second hydrothermal treatment corresponding to a second severity factory which is derived by treating the presoaked ligno-cellulosic biomass stream at at least a second hydrothermal temperature for at least a_second hydrothermal time, wherein the second severity factor is greater than the first severity factor.

It is also disclosed that the second hydrothermal treatment may conducted at at least a temperature in a range selected from the group consisting of 100°C to 220°C, 101°C to 220°C, 1 10° to 190°, and 130°C to 170°C.

It is further disclosed that the second hydrothermal treatment is conducted for at least a time in a range selected from the group consisting of 1 minute to 24 hours, 10 minutes to 12 hours, 15 minutes to 6 hours, and 20 minutes to 2 hours. BRIEF DESCRIPTION OF DRAWINGS

Figure 1 is a graph showing the amount of base added to neutralize the liquid sugar streams obtained by decationization hydrolysis, relative to the amount of sugar, according to one embodiment of the invention.

Figure 2 is a graph showing the percent dry matter of the chromatographic fraction obtained by separating the liquid sugar streams produced by decationization hydrolysis, according to one embodiment of the invention.

Figure 3 is a graph showing the percent dry matter of the chromatographic fraction obtained from the liquid sugar streams produced by enzymatic hydrolysis, according to another embodiment of the invention.

DETAILED DESCRIPTION

It is disclosed a process for producing a purified liquid stream from a ligno-cellulosic biomass. The purified liquid stream may be further refined to produce a refined liquid sugar stream. Both the streams are particularly suitable for feeding catalytical and biological conversion processes to produce biofuels and chemicals.

A description of a ligno-cellulosic feedstock may be found in WO2013098789, pag. 5- 10.

For the scope of the disclosed process, the ligno-cellulosic biomass comprises a ligno-cellulosic component and a non ligno-cellulosic water soluble compounds. The ligno-cellulosic component comprises carbohydrates (mainly glucans and xylans) and lignin, which may be then converted to biofuels and biochemicals. Carbohydrates are insoluble polymers of water soluble monomeric sugars (such as glucose and xylose).

The non ligno-cellulosic water soluble compounds comprise compounds different from carbohydrates and which are naturally present in the ligno-cellulosic biomass, including, among others, organic and inorganic water soluble salts of cations and anions including sodium, calcium, potassium, ammonium, magnesium; waxes and extractives in general. When solubilized in water, water soluble species are derived from these compounds by direct solubilization or also by more complex reactions.

Water soluble compounds are defined as follow: an amount of 50 g of ligno-cellulosic biomass is dispersed in 250 mL of distilled water at 65°C and shaked for 5 minutes. The slurry is filtered with a colander and the liquid fraction is collected and analyzed. Water soluble compounds are the compounds in the liquid fraction having a concentration greater than Og/1.

The ligno-cellulosic biomass may further comprise non ligno-cellulosic water insoluble compounds, such as intrinsic silica present in the ligno- cellulosic biomass, which are not solubilized in water at the conditions of the disclosed process.

The ligno-cellulosic biomass stream may further comprise water insoluble contaminants comprising for instance stones, gravel, sands, sand, dust, clay, silica and silicates in general, and metal objects, which are collected with the ligno-cellulosic biomass in harvesting and handling operation of the ligno-cellulosic biomass and it is desirable that they are separated from the ligno-cellulosic biomass before feeding the ligno-cellulosic biomass to downstream devices, which could be damaged.

The size of water insoluble contaminants may vary from very small particles, in the sub millimeter range as in the case of sand, to many centimeters, as in the case of stones. They are in general mixed with the ligno-cellulosic biomass and may adhere on the surface of the ligno- cellulosic biomass or be present in bundles of the ligno-cellulosic biomass. In these cases, separation from the ligno-cellulosic biomass may be difficult.

Preferably, the percent amount of the contaminants in the ligno- cellulosic biomass stream is less than a 10%, 5%, 3%, and 1% by weight of the ligno-cellulosic biomass stream.

In the disclosed process, the ligno-cellulosic biomass is subjected to a first hydro thermal treatment in water or in a liquid comprising water to solubilize at least a portion of the non ligno-cellulosic water soluble compounds.

The first hydrothermal treatment produces a pre-soaking mixture comprising a pre-soaking liquid and a pre-soaked ligno-cellulosic biomass. The first hydrothermal treatment is a pre-soaking of the ligno- cellulosic biomass, as in a preferred embodiment there is a following soaking step of the pre-soaked ligno-cellulosic biomass.

The pre-soaking liquid comprises water and water soluble species derived from the non ligno-cellulosic water soluble compounds of the ligno-cellulosic biomass. The water soluble species comprises in particular charged species derived from the solubilization of a first portion of water soluble salts in the ligno-cellulosic biomass. The pre- soaking liquid contains soluble compounds which are detrimental for the following steps of the disclosed process, and it is withdrawn from the following process steps.

As known in the art, the process conditions of a hydrothermal treatment may be combined to express a severity factor of the treatment. The severity factor Ro is expressed by the formula Ro= t x exp[(T- To/ 14.75)], wherein t is the hydrothermal time in minutes, T is the hydrothermal temperature in °C, and To= 100°C.

In the case than the hydrothermal treatment is conducted in acid conditions, for instance by adding a mineral or organic acid to lower the pH, the severity factor Ro my be expressed by the formula Ro= t x exp[(T-To/ 14.75)]-pH. In the case that the hydrothermal treatment comprises more than one hydrothermal steps, being each step characterized by a temperature, a time and optionally a pH, the severity factor of each step may be defined according to the previous formulas, and the severity factor of the hydrothermal treatment may be expressed as the sum of the severity factors of each step. The total severity factor may be expressed also as log(Ro).

In one embodiment, the first hydrothermal treatment comprises a single hydrothermal step, that is it conducted at a first hydrothermal temperature for a first hydrothermal time, corresponding to a first severity factor. In a preferred embodiment, the first hydrothermal treatment comprises two hydrothermal steps, which may be conducted at different conditions of temperature and time. The first severity factor may be expressed as the sum of the severity factors of the two steps.

During the first hydrothermal treatment the ligno-cellulosic biomass is preferably agitate, or mixed, by means of mechanical means to promote the solubilization of non ligno-cellulosic water soluble compounds. In another preferred embodiment, the pre-soaked ligno-cellulosic biomass is compressed preferably by means of a compression screw to remove impregnated liquid. The compression may be done after the first hydrothermal treatment, or during the first hydrothermal treatment, or between two steps of the first hydrothermal treatment. The pre-soaked ligno-cellulosic biomass is therefore separated from the pre-soaking mixture and further subjected to a treatment to produce a mixture of a liquid biomass feed stream and a solid treated ligno- cellulosic biomass stream. The solid treated ligno-cellulosic biomass is separated from the mixture and it may then be further treated for feeding another conversion process or processes to different chemical end-products, while the liquid biomass feed stream is further processed to produce the liquid sugar stream according to the disclosed process.

The liquid biomass feedstream comprises water and water soluble oligomeric sugars, which are derived from the solubilization of at least a portion of the hemicellulose and cellulose. Even if any treatment may be used to produce the liquid biomass feedstream, a preferred treatment comprises a second hydrothermal process, which solubilizes a portion of the carbohydrates in the feedstock without the use of any chemicals or catalysts. Preferably, the second hydrothermal treatment comprises a soaking step of the pre-soaked ligno-cellulosic feedstock to produce a soaked liquid stream comprising soluble sugars, mainly derived from the hemicellulose of the ligno-cellulosic biomass.

The second hydrothermal treatment may comprise more than one step.

In a preferred embodiment, the second hydrothermal treatment is conducted at process conditions corresponding to a second severity factor, and the second severity factor is greater than the first severity factor.

In another embodiment, the second hydrothermal treatment is conducted at a second hydrothermal temperature which is greater than the first hydrothermal temperature. In the case that the first hydrothermal treatment comprises more than one step, being conducted at different temperatures, the second hydrothermal temperature is preferably greater than the maximum temperature occurring during the first hydrothermal treatment. In the case that the second hydrothermal treatment comprises more than one step, conducted at different temperatures, the maximum temperature occurring during the second hydrothermal treatment is preferably greater than the maximum temperature occurring during the first hydrothermal treatment. The first hydrothermal treatment may be conducted at a temperature in the range of greater than 100°C to 150°C, according to the teaching of WO2013098789, which is herein incorporated by reference. Pressures greater than lbar are involved and thereby pressure vessels, special heating kits, insulation may be required, increasing the capital and operating cost of the process. Thereby, in a preferred embodiment, the first hydrothermal treatment is be conducted at a temperature in a range of 30°C to 100°C, preferably of 40°C to 99°C, more preferably of 40°C to 90°C, and most preferably of 50°C to 85°C. In the case that the first hydrothermal treatment comprises more than one hydrothermal step, the same preferred temperature ranges apply to each hydrothermal step.

The first hydrothermal treatment may be conducted for a time in a range selected of 10 seconds to 300 minutes, preferably of 1 minute to 20 minutes, more preferably of 2 minutes to 20 minutes, even more preferably of 2 minutes to 15 minutes, and most preferably of 3 to 10 minutes. In the case that the first hydrothermal treatment comprises more than one hydrothermal step, the same preferred time ranges apply to each hydrothermal step.

The second hydrothermal treatment may be conducted at a temperature in a range of 100°C to 220°C, preferably of 101°C to 220°C, more preferably of 1 10° to 190°, and most preferably of 130°C to 170°C. In the case that the second hydrothermal treatment comprises more than one hydrothermal step, the same preferred temperature ranges apply to each hydrothermal step.

The second hydrothermal treatment may be conducted for a time in a range of 1 minute to 24 hours, more preferably of 10 minutes to 12 hours, even more preferably of 15 minutes to 6 hours, and most preferably of 20 minutes to 2 hours. In the case that the second hydrothermal treatment comprises more than one hydrothermal step, the same preferred time ranges apply to each hydrothermal step.

The preferred processes for producing the liquid biomass feedstream from the ligno-cellulosic biomass is described in details in the following of the present specification.

The liquid biomass feedstream comprises thereby water and water soluble oligomeric sugars. The water soluble oligomeric sugars may comprise any sugar soluble in water different from monomeric sugars. Thereby, water soluble oligomeric sugars may comprise C5 and C6 oligomeric sugars. C5 sugars are pentose-based sugars, wherein pentose is a monosaccharide with five carbon atoms. Xylose is an example of monomeric pentose sugar. C6 sugars are hexose-based sugars, being the hexose a monosaccharide with six carbon atoms. Glucose is an example of monomeric hexose sugar. In a preferred embodiment the water soluble oligomers comprises mainly xylooligomers, which are derived from the hemicellulose of the ligno- cellulosic biomass. C5 sugars rich-streams represent a low-value stream which may be derived from the ligno-cellulosic biomass. As examples, C5 sugars rich-streams may be produced also as a byproduct in the pulp and paper industry. Besides oligomeric sugars, the liquid biomass feed stream comprises soluble salts which are disassociated into anions and cations. These salts, which may be organic and inorganic salts, are the salts of cations including sodium, calcium, potassium, ammonium, magnesium, and others cations contained in the original ligno-cellulosic biomass and solubilized by the treatment of the ligno-cellulosic biomass.

The liquid biomass feed stream may further comprise other soluble compounds, including monomeric sugars such as glucose, xylose, arabinose, galactose, mannose; acetic acid, lactic acid, oxalic acid, among other organic acids, and the salts of these acids. A variety of other compounds may be present in the liquid biomass feed stream, including sugar degradation products such as furfural and hydroxymethyl furfural, and soluble phenolic compounds derived from lignin. Organic extractive compounds, such as soaps and fatty acids, may also be present. Preferably, the liquid biomass feed stream does not comprise anions from added inorganic acids.

In a preferred embodiment, the liquid biomass feedstream is subjected to a removal step of suspended solids, as suspended solids may obstruct downstream equipments. Removal of suspended solids may include for instance, but is not limited to, the use of a press, a decanter, a centrifuge, a filter, a flocculating agent, a microfilter, a plate and frame filter, a crossflow filter, a pressure filter, a vacuum filter, or a combination thereof.

In an embodiment, the liquid biomass feedstream is subjected to a concentration step, which may be carried out using any technique known to those of skill in the art. For example, concentration may be carried out by subjecting the liquid biomass feedstream to membrane filtration, evaporation, or a combination thereof. Without being limiting, microfiltration (with a pore size of 0.05 to 5 microns) may be carried out to remove particles, followed by ultrafiltration (500-2000 raw cut off) to remove soluble lignin and other large molecules and reverse osmosis to increase the concentration of soluble compounds, followed by evaporation.

Preferably, the liquid biomass feedstream has a dry matter content by weight which is greater than 3%, more preferably greater than 5% even more preferably greater than 10%, even yet more preferably greater than 15%, and most preferably greater than 20%.

The total percent amount of soluble sugars in the liquid biomass feed stream, which include oligomeric sugars and, if present, monomeric sugar, is preferably greater than 30%, 40%, 50%, 60%, 70%, 80% by weight on a dry basis. For instance, if the liquid biomass feed stream has a dry matter content of 10% and the total amount of soluble sugars is 50%, there are 50grams of soluble sugars in lKg of liquid biomass feed stream.

The total percent amount of soluble salts in the liquid biomass feed stream may be 10% to 70% by weight on a dry basis, or 10% to 60%, or 10% to 50%, or 10% to 40%, or 10% to 30%.

According to the disclosed process, the liquid biomass feed stream is then subjected to an hydrolysis step, which is conducted at hydrolysis conditions sufficient to hydrolyze at least a portion of the oligomeric sugars of the liquid biomass feed stream to monomeric sugars. The hydrolysis is conducted at a pH which is less than 5.5, and an acid or a base may be added to reach the desired pH value.

In one embodiment, the hydrolysis is an enzymatic hydrolysis and it is conducted by contacting the liquid biomass feed stream with an enzyme or enzyme cocktail. Preferably the pH of the liquid biomass feed stream has to be adjusted to a value in the range between 4.0 and 5.5, and more preferably between 4.5 and 5.5, prior to contact the liquid sugar stream with the enzyme or enzyme cocktail.

The enzyme composition can comprise any protein useful in degrading or converting the oligomeric sugars of the liquid biomass feed stream. In a preferred embodiment, the enzyme cocktail comprises one or more (e.g., several) hemicellulolytic enzymes. The term "hemicellulolytic enzyme" or "hemicellulase" means one or more enzymes that hydrolyze a hemicellulosic material. The hemicellulase is preferably one or more enzymes selected from the group consisting of an acetylmannan esterase, an acetylxylan esterase, an arabinanase, an arabinofuranosidase, a coumaric acid esterase, a feruloyl esterase, a galactosidase, a glucuronidase, a glucuronoyl esterase, a mannanase, a mannosidase, a xylanase, and a xylosidase. Preferably, the hemicellulase is selected from the group of xylanases and xylosidases. The term "xylanase" means a 1 ,4-beta-D-xylan- xylohydrolase (E.C. 3.2.1.8) that catalyzes the endohydrolysis of 1 ,4- beta-D-xylosidic linkages in xylans.

The term "beta-xylosidase" means a beta-D-xyloside xylohydrolase (E.C. 3.2.1.37) that catalyzes the exo-hydrolysis of short beta (1→4)- xylooligosaccharides, to remove successive D-xylose residues from the non-reducing terminals.

In one embodiment, the enzyme cocktail further comprises one or more proteins/ polypeptides selected from the group consisting of a cellulase and a hemicellulase. The cellulase is preferably one or more (e.g., several) enzymes selected from the group consisting of an endoglucanase, a cellobiohydrolase, and a beta-glucosidase.

Enzymatic hydrolysis is preferably carried out in a suitable aqueous environment under conditions that can be readily determined by one skilled in the art. In one aspect, hydrolysis is performed under conditions suitable for the activity of the enzymes, i.e., optimal for the enzymes. The hydrolysis can be carried out as a batch, a fed batch or continuous process where the liquid biomass feed stream is fed gradually to, for example, an enzyme containing hydrolysis solution. The enzymatic hydrolysis is generally performed in a vessel, such as for example a stirred-tank reactor under controlled pH, temperature, and mixing conditions. Suitable process time, temperature and pH conditions can readily be determined by one skilled in the art. For example, the enzymatic hydrolysis can last up to 200 hours, but is typically performed for preferably about 6 to about 120 hours, e.g., about 16 to about 72 hours or about 24 to about 48 hours. The same time ranges apply to residence time in the case of continuous hydrolysis. The temperature is in the range of preferably about 25° C to about 70° C, e.g., about 30° C to about 65° C, about 40° C to about 60° C or about 50° C to about 55° C.

In a preferred embodiment, the hydrolysis of the liquid biomass feed stream comprises at least two steps. The first step is to create an acidic stream from the liquid biomass feed stream having a pH which is less than 3, more preferably less than 2.5, even more preferably less than 2, and most preferably less than 1. This is accomplished by increasing the amount of H ⁺ ions to the liquid biomass feed stream to create the acidic stream.

After the desired pH is obtained, the next step is hydrolyzing the oligosaccharides in the acidic stream by raising the temperature of the acidic stream to a hydrolysis temperature in a range between 80°C and 200°C for the hydrolysis reaction to occur, creating a hydrolyzed stream. After hydrolysis, the hydrolyzed stream can be passed to other unit operations for further processing.

In one embodiment, the acidic stream is created by adding a mineral acid to the liquid biomass feed stream. Several types of acids, concentrated or diluted, can be used, such as sulfurous, sulfuric, hydrochloric, hydrofluoric, phosphoric, nitric and formic acid. Sulfuric and hydrochloric acids are the most commonly used catalysts for hydrolysis of ligno-cellulosic biomass. The acid concentration used in the concentrated acid hydrolysis process is in the range of 10-30%. The process occurs at low temperatures, producing high hydrolysis yields of monomers (i.e. 90% of theoretical glucose yield). In a preferred embodiment, the acidic stream is created according to the teaching of WO2013026849, which is incorporated herein by reference. The disclosed process takes advantage of the salt content of the liquid biomass feed stream. In order to obtain the required acidity for the hydrolysis step, the content of salts in the feed stream is reduced via cation exchange while at the same time replacing the cations with H ⁺ ions. While the salts may naturally occur in the biomass, they can also be added as part of the pre-treatment processes or prior to or during the creation of the acidic stream. The process of reducing the amount of cations of the soluble salts in the liquid biomass feed stream, called decationization, removes the cations by exchanging them with H ⁺ ions. One way the cations in the liquid biomass feed stream can be replaced by H ⁺ ions is by using an ion exchange agent, such as an ion exchange resin. The cations can also be exchanged using a membrane. For example, Dupont's Nafion(R) PFSA Resins can be used as resins in an exchange column or as a membrane through which the solution is passed. These are per- fluorinated resins in the sulfonyl fluoride (-SO2F) form. Preferably, the ion exchange agent is a strong cationic resin, more preferably a sulphonated resin in Hydrogen form, such as for instance Relite EXC14 by Mitsubishi Chemical Corporation, Japan.

An additional step of separating at least a portion of the acidic stream from the ion exchange media before subjecting the separated portion to the hydrolysis temperatures may be needed. Preferably, all the ion exchange media is removed from the acidic stream before hydro lyzing the oligosaccharides in the acidic stream.

It should be recognized that the amount of salts present influences the amount of H ⁺ ions that can be increased (added to the liquid) via ion exchange. The amount of H ⁺ ions also determines the pH of the acidic stream. These salts can be concentrated according to the steps outlined above. Should the feed stream not have sufficient salts with cations, one can add a salt or cations in another manner to the liquid biomass feed stream prior to, and/ or during, and/ or after the creation of the acidic stream, or combination thereof. Preferably, the salts of Magnesium, Calcium, Sodium, Potassium can be used. Preferably salts with a monovalent cation are used as the cation will not reduce the efficiency of the ion exchange media as much as a bivalent ion.

In the disclosed process, the majority of the H ⁺ ions are derived from the decationization of the liquid biomass feed stream by means of the ion exchange agent. Preferably, more than 60%, 70%, 80%, 90% and 95% of the H+ ions come from the decationization step. In a preferred embodiment, 100% of the H ⁺ ions comes from the decationization step, thereby no H ⁺ ion comes from added acids such as sulfuric acid and no additional anion is further added to the biomass feedstream.

If the cations in the liquid biomass feed stream are not in a sufficient amount to reach the desired pH value, or if it is desired to remove only a portion of the cations, additional H ⁺ ions may be added to the stream. The amount of H ⁺ ions can be increased via any known means, including the use of acids, electrical current, the addition of hydrogen peroxide, and the use of a membrane; or even in- situ production of the H ⁺ ions. Increasing the amount of H ⁺ ions, or protons, in-situ can be accomplished by adding a compound which does not contain H ⁺ ions capable of disassociating in water, but rather catalyzes a reaction, or the compound itself reacts, with component(s) already present in the liquid biomass feed stream. For example, AlC contains no H ⁺ ions. However, when added to the liquid biomass feed stream, the AlCb will react with the water in the liquid biomass feed stream to form Al(OH)3 and HC1, thus creating the H ⁺ ion. In this manner, the amount of the H ⁺ ions are increased without adding H ⁺ ions to the liquid biomass feed stream. The percent of the total amount of H ⁺ ions derived from an acid or acids added to the liquid biomass feed stream may be of less than 50%, less than 40%, less than 30%, less than 20%, less than 10%, and 1%.

In the case of decationization, the pH of the decationized stream becomes lower than the pH of the feedstream. The pH that can be achieved with decationization depends on the initial cation concentration in the feed liquid, the cations added to the stream, the anions present in the stream, the ion resin exchange capacity, specific velocity through the resin and temperature of exposure. Therefore, the decationization should occur at a temperature in the range of 5°C to 60°C, and for a time sufficient to reach a pH of the liquid biomass feed stream which is less than 3.0, more preferably less than 2.5, more preferably less than 2.0, more preferably less than 1.5. One of ordinary skill knows that pH has a lower theoretical limit of up to but not including 0, thus each of the above numbers can be expressed as the upper limit of the pH of the acidic stream, with the pH being greater than, but not including, 0.0.

Once the desired pH is reached, the acidic stream is hydrolyzed by increasing the temperature of the acidic stream to a hydrolysis temperature greater than 80 °C, and preferably within the range of 80 °C to 200 °C. Other ranges are 80 °C to 180 °C, 95 °C to 180°C, 100 °C to 180 °C, 120 °C to 180 °C and 120 °C to 170 °C the most preferred. The hydrolysis temperature is maintained for a time sufficient to hydrolyze the components (oligosaccharides) to the degree desired. The hydrolysis time may be from 1 second to 24 hours, or from 30 second to 18 hours, or from 1 minute to 12 hours. Hydrolysis temperature and time are preferably selected to significantly reduce the formation of degradation products. As a result of the decationization and hydrolysis steps, the large amounts of acid or acids used in the hydrolysis step is avoided, allowing the passage from a rather harsh treatment to a totally mild one and the consumption of acid used can be reduced to the amount needed to regenerate the cationic resin (or not used at all). The acid is then recovered in a separate stream and then more easily disposed of.

At the same time, the acidic liquid biomass feed stream after hydrolysis is a cleaner liquid, comprising monomeric sugars, low content of salts and low amount of degradation products that could hinder subsequent chemical or biological transformations of the sugars.

Inventors have found that increasing the amount of H ⁺ ions and removing cations at the same time greatly improves, or simplifies, the downstream purification step or steps, permitting to recover a purified liquid sugar stream from the liquid biomass feed stream with a high recovery yield.

After hydrolysis, the pH of the acidic liquid biomass feed stream is increased to a value compatible with the following purification step which comprises the use of a cationic resin. Thereby, the pH of the liquid biomass feed stream is preferably raised to a value in the range of 5.5 to 8, more preferably of 5.5 to 7, even more preferably of 5.5 to 6.5, and most preferably of 6 to 6.5. The increase of pH to the desired value is obtained by adding an amount of a reference base comprised of a reference cation and a reference anion.

The term "base" is meant to encompass any soluble species that, when added to pure water, gives a solution with a pH that is more than 7. Preferably, the pH adjustment is carried out using a soluble base. By the term "soluble base", it is meant a base that has a solubility in water that is at least 0.1 M at 20° C. This term is meant to exclude salts that are slightly soluble or insoluble. Examples of insoluble bases that are excluded from the definition of soluble base are CaCO3 and, Ca(OH)2. Non-limiting examples of soluble bases include sodium hydroxide, potassium hydroxide, ammonia, and ammonium hydroxide. In the liquid sugar stream after neutralization a huge predominance of only one species of cation, namely the reference cation, is obtained, starting from a plurality of cations of different species in the liquid biomass feed stream, which may be monovalent and multivalent. The liquid sugar stream after neutralization comprises water, monomeric sugars from the hydrolysis step and a salt or salts comprised of the reference cation and the anions of the liquid biomass feed stream. This concept can be conveniently expressed by the ratio of the total ionic equivalents of cations different from the reference cations to the total ionic equivalents of cations in the liquid sugar stream, which is less than 0.20, preferably less than 0.15, more preferably less than 0.10, even more preferably less than 0.05 and most preferably less than 0.03. Ionic equivalents are introduced because cations with different valence may affect differently the cationic resin of the following purification step. The equivalent, sometimes termed the molar equivalent, is a unit of electrical charge used in chemistry. It could be converted to Coulombs 'C (SI unit) using Faraday's constant F in 'C/mol', where 1 'eq' = F 'C.

The equivalent of substance A is the amount of substance A (in moles) multiplied by its valence. The equivalent could be also formally defined through the amount of substance which will either react with or supply one mole of hydrogen ions (H ⁺ ) in an acid-base reaction; or react with or supply one mole of electrons in a redox reaction.

Practically, certain amount of univalent ions provides the same amount of ionic equivalents while the same amount of divalent ions provides twice the amount of ionic equivalents. For example, 1 mmol of Na ⁺ is equal 1 mEq, while 1 mmol of Ca ⁺⁺ is equal 2 mEqs. In an embodiment, the soluble reference salt comprised of the reference cation and the dissociated anion of the liquid biomass feed stream may be at least partially precipitated in solid form. Precipitation may be driven by pH change, concentration of the soluble reference salt, temperature as examples. Precipitation is preferably conduced before subsequent purification step in such a way that precipitated compounds do not enter purification equipment.

The liquid sugar stream is then treated by a cationic exchange resin to separate sugars and other nonionic compounds from the salts and other ionic compounds, thereby producing a purified liquid sugar stream comprising the major portion of the monomeric sugars, and at least a second stream, comprising at least the major portion of the salts and other ionic compounds. In the disclosed process, the cation of the cationic exchange resin is the same species of the reference cation of the reference base added to raise the pH of the acidic biomass feed stream, as the presence of the same species of cations in the liquid sugar stream and the cation resin improves the separation yield, the purity of the product streams and the lifetime of the resin, that is the time that the resin can work with high efficiency before being regenerated.

Thereby, the purified liquid sugar stream is characterized by a purity ratio, which the total amount of monomeric sugars to the total amount of all the compounds on a dry basis, and the purity ratio of the purified liquid sugar stream is greater than 60%, preferably greater than 70%, more preferably greater than 80%, even more preferably greater than 85%, and most preferably greater than 90%. The separation process is further characterized by a recovery ratio which is the ratio of the total amount of monomeric sugars in the purified liquid stream to the total amount of monomeric sugars in the liquid sugars stream. The recovery ratio may be greater than 60%, preferably greater than 70%, more preferably greater than 80%, even more preferably greater than 85%, even yet more preferably greater than 90%, and most preferably greater than 95%.

Preferably, the cationic resin is a strong cation resin, for example, which is not to be considered limiting, with a polystyrene backbone and divinylbenzene crosslinking. These resins have sulfonate functional groups and are available commercially in the sodium form, or, less preferably, in the hydrogen, potassium or ammonium form. The resins are preferably of diameter of from about 0.1 to about 1.0 mm. Cationic exchange resins are available from several vendors, including Dow or Mitsubishi. Preferably, the reference cation is a monovalent cation, such as Sodium or Potassium. More preferably, the reference cation is Na ⁺ and the preferred base is Sodium hydroxide. Thereby, the cationic resin used for separating the liquid sugar stream is a resin in sodium form. In the case of hydrolysis by decationization step, the removal of the major portion of divalent ions from the liquid biomass feed stream occurs in the decationization step. Divalent cations, such as Magnesium and Calcium, are known to poison the monovalent cationic resin very quickly, thereby reducing the separation efficiency of the resin. Thereby the removal of divalent cations occurs without any devoted softening step or equipment. In the case of enzymatic or acid hydrolysis by adding a mineral acid, a softening step is instead desirable. Being the cation of the cationic resin a monovalent cation, the liquid sugar stream may be further characterized by the ratio of the total ionic equivalents of non- monovalent cations to the total ionic equivalents of cations, which may be less than 0.05, preferably less than 0.03, more preferably less than 0.02 and most preferably less than 0.01. The separation of the liquid sugar stream in the product streams may occur mainly by means of ion-exclusion principle. The ion exclusion system of the present invention may be operated in a temperature range of about 20° C to about 90° C, preferably at a temperature between about 45° C to about 80° C. The process of ion exclusion separation may involve the use of one, or more than one, column filled with ion exchange resin, as it is evident to one of skill in the art. For the sake of simplicity, the operation of a single column will be illustrative, but the use of more than one column is also considered to be within the scope of the present invention. The column may be prepared prior to carrying out the separation by converting it into the desired cationic form, which is preferably Sodium form. This may involve washing a volume of the liquid sugar stream through the column. Alternatively, the column may be prepared by washing it with a volume of solution containing cations corresponding to the reference cation of the liquid sugar stream. Once the column is in the appropriate cation-exchange form, the liquid sugar stream is applied onto the column. A desired liquid flow rate is selected as may be readily determined by one of skill in the art, for example, but not limited to, a liquid flow rate corresponding to about 5% to about 70% of the column volume per hour. As the purified sugar stream is applied, the charged ions in the salts and other charged compounds are excluded from the resin and flow through the column. The sugar and other nonionic compounds are not repelled by the charged resin, and penetrate the pores of the resin. The sugar and other nonionic compounds are thereby retained by the resin and elute the column more slowly than the ionic compounds. After the desired volume of the clarified sugar stream is injected, the feed is switched to water. The ionic compounds flow through the column and are collected in one or more than one by-product stream. This one or more than one by-product stream contains the majority of the reference salt and other ionic compounds, and small amounts of sugar and it is followed by the elution of sugars arising from the processing of the cellulosic biomass and nonionic compounds, which are collected separately from the one or more than one by-product stream. The purified liquid sugar stream contains most of the sugar and a few amount of the salt and other ionic and neutral components.

As a result of the decationization and hydrolysis steps, inventors have found that the separation of the purified liquid sugar stream and the second stream from the liquid sugar stream is greatly improved with respect to the case of acid hydrolysis. It is believed that this is due to the lower content of salts and other charged species in the liquid sugar stream.

Thereby, by the disclosed process it is possible to recover a highly pure sugar stream wasting few amount of sugars of the feed stream.

In a preferred embodiment, the ion exclusion chromatography is carried out by a Simulated Moving Bed (SMB) device. An SMB contains ion exchange resin similar to that in an ion exclusion system described above, and performs the same type of separation of sugars and nonionic compounds in the purified liquid sugar stream and salts and other ionic compounds in the second stream. For a given feed stream, an SMB is usually run at the same pH and temperature as an ion exclusion system. However, an SMB system has distinct locations for feeding of the liquid sugar stream, feeding of dilution water, and withdrawal of purified liquid sugar stream and second stream. For example, which is not to be considered limiting, four flow locations equally spaced apart may be used on one or more than one column. In a preferred embodiment, one column is used. This simplifies the demarcation of zones and allows for a given column to be brought off line, for cleaning or maintenance without overly disturbing the operation. For example, which is not to be considered limiting, from about 4 to about 16 columns may be used. In a more preferred embodiment, about 4 to about 8 columns are used. However, the number of columns may be adjusted as required.

Another difference between an SMB and a single-column ion exclusion system is that the SMB has a recirculation flow that supplements and is co-current with all of the other flow streams. This recirculation flow is carefully chosen, along with the other flows, to provide the optimum separation between the sugar and salt streams. Additionally, an SMB system simulates movement of the resin bed in a direction opposite to that of the liquid flow.

Improved SMB ("ISMB") systems may also be used as described herein. ISMB systems include variable flow rates of feed, dilution water, product, raffinate, or a combination thereof, or sequential periods with one or more streams closed off, with or without re-circulation of the liquid in the columns, or a combination of two or more of these features. The present invention can be practiced with ISMB or SMB operations, or with Sequential Simulated Moving Bed.

A detailed description on how to operate a SMB separation system may be found in Jari Heinonen, Tuomo Sainio, "Chromatographic Fractionation of Lignocellulosic Hydrolysates", Advances in Chemical Engineering, Volume 42, pag.261.

In a preferred embodiment, the cationic resin is selected to have also a size-exclusion action, beside a ion-exclusion action. The second stream separated from the liquid sugar stream will then comprise also a portion of non-ionic compounds present in the liquid sugar stream. As size exclusion is known in the art, a person skilled in the art knows how to select the cation resin to operate the separation also on the basis of size exclusion principle. The purified liquid sugar stream may be further subjected to a refining step to further remove compound which are not monomeric sugars from the purified liquid sugar stream. These compounds are both charged species and neutral compounds present in the liquid sugar stream which have been retained purified liquid sugar stream in the ion- exclusion separation step.

By means of the refining step, an ultra-pure refined liquid sugar stream may be obtained, having a purity ratio greater than 90%, 95%, 97%, 98%, 99% and 99.9%. The purity ratio of the refined liquid sugar stream is percent ratio of the total amount of monomeric sugars to the total amount of all the compounds on a dry basis in the refined liquid sugar stream.

Moreover, the recovery of the refined liquid sugar stream may occur with great efficiency and very few sugars may be lost in the refining step. The recovery ratio of the refining step, which is the ratio of the total amount of monomeric sugars in the refined liquid stream to the total amount of monomeric sugars in the purified liquid sugars stream entering the refining step, may be greater than 95%, preferably greater than 97%, more preferably greater than 98% and most preferably greater than 99%. In a preferred embodiment the removal of the compounds different from monomelic sugars is conducted by means of at least a refining material, and removal occurs by adsorbing the compounds on the refining material. The purified liquid sugar stream is contacted with the refining material or materials, preferably in a column configuration, without adding any further dilution liquid as in the case of chromatographic separation. As the purified liquid sugar stream of the disclosed process contains already a limited amount of compounds different from monomeric sugars, regeneration of the refining material may not be needed often, being adsorption mechanism compatible with an industrial process.

The refining material may be an ion-exchange resin, thereby the removal occurs by means of ion exchange. The ion exchange resin is preferably a strong anion-exchange resin in OH ^" form, which retains at least a portion of the anions in the purified liquid sugar stream on positively charged functional groups. In the process, anions are exchanged with OH ^" hydroxyl groups. Thereby, the pH of the refined liquid sugar stream is preferably raised to a value in the range of 3 to 12, more preferably of 6 to 12, even more preferably of 6 to 1 1 , and most preferably of 7 to 10. The removal by means of ion exchange should occur at a temperature in the range of 5°C to 120°C, or 5°C to 100°C, or 5°C to 80°C, or 5°C to 70°C _o or 5°C to 60°C, or 20°C to 60°C.

Even if any kind of strong anionic resin may be used, a copolymer styrene-DVB anion exchange resin containing strongly basic ions groups is preferred. An example of such a resin is Relite EXA268 from Mitsubishi Chemicals, Japan.

In another embodiment, the ion exchange resin is a strong cation- exchange resin in H ⁺ form, adsorbing at least a portion of the remaining cations in the purified liquid sugar stream. This resin is preferably similar to the resin used in the decationization step. The refining material may be a carbonaceous material, preferably an activated carbon. In a preferred embodiment, the refinement of the purified liquid sugar stream comprises two refining materials. The purified liquid sugar stream is first contacted with an ion-exchange resin and then with an activated carbon. The two refining materials are placed in separated columns and the purified liquid sugar stream is inserted into the first column containing the ion-exchange resin, which can be an anionic exchange resin or a cationic exchange resin or a mixture of the two, to remove remnant charged species. The purified liquid sugar stream, after first refinement, is extracted from the first column and inserted into the second column containing the activated carbon to remove preferentially non charged species. The stream exiting from the second column is the refined liquid sugar stream.

In a preferred embodiment, the liquid biomass feed stream is a C5 sugars rich-streams derived mainly from the solubilization of the hemicellulose of the ligno-cellulosic biomass and comprises mainly xylooligomers. Thereby, the refined liquid sugar streams comprises mainly xylose and other C5 monomeric sugars, such as arabinose. Preferably, the percent amount of C5monomeric sugars on the total amount of monomeric sugars in the refined liquid sugar stream is greater 50%, more preferably greater than 60%, even more preferably greater than 70%, even yet more preferably greater than 80%, and most preferably greater than 90%.

The purified liquid sugar stream or the refined liquid sugar stream may be converted by means of a biological or catalytic conversion process to biofuels or chemicals.

In a preferred embodiment, the refined liquid sugar stream is converted to a polyols mixture comprising ethylene glycol by means of a catalytic conversion process comprising a hydrogenation step of the refined liquid sugar stream to produce a hydrogenated mixture comprising water and a mixture of sugar alcohols; and a hydrogenolysis step of the hydrogenated mixture, to produce the polyols mixture. In the hydrogenation step, the refined liquid sugar stream is contacted with a hydrogenation catalyst and hydrogen at hydrogenating conditions promoting the hydrogenation of the sugars in the liquid sugar stream. Preferred sugar alcohols are xylitol, sorbitol and arabitol, or mixture thereof. Even more preferably, the hydrogenated mixture comprises xylitol and the preferred amount of xylitol in the hydrogenated mixture on a dry basis is greater than 45%, more preferably greater than 70%, even more preferably greater than 80%, yet even more preferably greater than 90%, being greater than 95% the most preferred value.

The hydrogenolysis reaction of the sugar alcohols in the hydrogenated mixture produces a hydrogenolysis mixture comprising water, ethylene glycol, and propylene glycol. It may further comprise glycerol and other polyols, unwanted compounds, comprising acid lactic or formic acid, and unreacted sugar alcohols.

The polyols stream comprising ethylene glycol and propylene glycol may be recovered from the hydrogenolysis mixture by any process known in the art and still to be invented.

An ethylene glycol stream and a propylene glycol stream may then be separated from the polyols mixture preferably by means of distillation. The ethylene glycol stream comprises a plurality of diols, wherein ethylene glycol is the main component, as the amount of ethylene glycol, expressed as molar percent with respect to the plurality of diols, is preferably greater than 80%. In an embodiment, the ethylene glycol stream further comprises at least one diol selected from 1 ,2-Propylene glycol, 1 ,2-Butanediol and 1 ,2- Pentanediol.

The ethylene glycol stream may be used to produce a polyester resin. In a preferred embodiment, at least 85% of the acid moieties of the polyester are derived from terephthalic acid or its dimethyl ester.

In both method, the polyester may be further polymerized to a higher molecular weight by a solid state polymerization, which is particularly useful for container (bottle) application. Preferred first hydrothermal treatment of the ligno-cellulosic feedstock

In a preferred embodiment, the first hydrothermal treatment is a continuous process, wherein the ligno-cellulosic feedstock stream is introduced in an extraction vessel containing an extraction solution to solubilize at least a portion of the water soluble non-ligno-cellulosic compounds. The extraction solution comprises water and dissolved water soluble species derived from a previously treated aliquot of the ligno-cellulosic biomass. Stated in other words, the water soluble species derived from the ligno-cellulosic biomass stream are accumulated in the extraction solution.

In an embodiment, the ligno-cellulosic biomass is first treated in the extraction vessel in a great amount of the dirty extraction solution, having a high concentration of water soluble species, then extracted from the extraction vessel as a slurry. The majority by weight of free liquid is removed by draining, while optionally rinsing the pre-soaked ligno-cellulosic biomass with low flow of a clean rinse solution stream. The rinsing may be considered a second step of the hydrothermal treatment, as additional water soluble species may be solubilized in the rinsing step. In order for the hydrothermal treatment to be continuous, it is not necessary that the ligno-cellulosic biomass stream is continuously introduced into the extraction vessel, but it can be introduced at steady aliquots or pulses. Thus there are moments when there is no ligno- cellulosic biomass entering the extraction vessel. But, over time, the total mass introduced into the extraction vessel equals the total mass removed from the extraction vessel. One distinguishing feature between a continuous and a batch process is that, in a continuous process, the separation step is occurring or progressing at the same time that either the ligno-cellulosic biomass is introduced into the extraction vessel and/ or the pre-soaked ligno-cellulosic biomass is removed from the extraction vessel. Another way to state this is that the hydrothermal treatment in the extraction vessel occurs while simultaneously, or at the same time, removing the pre-soaked ligno-cellulosic biomass from the extraction vessel. Such removal is done in a continuous manner which includes an aliquot or pulse removal.

The extraction vessel may be of any of size and shape suitable for the scope of the disclosed process. The extraction vessel may be an open vessel, with a free surface of the extraction solution exposed to the external environment, or a closed vessel, with a cover to insulate the extraction solution from the external environment.

Preferably, the extraction vessel has the shape of a pool, with an elongated horizontal section, with a main dimension, or length, which may be between 2m and 100m, preferably between 4m and 80m, even more preferably between 4m and 40m. In the case that the ligno-cellulosic biomass stream is compacted in bales, the bales are preferably disaggregated for introducing the loose ligno-cellulosic biomass stream into the extraction vessel.

The ligno-cellulosic biomass stream is preferably introduced into the extraction vessel as a dry biomass, meaning that no free liquid in present in the incoming stream. The moisture content is preferably less than 50%, more preferably less than 30%, even more preferably less than 20%, and most preferably less than 10%. In another embodiment, the ligno-cellulosic biomass stream is introduced into the extraction vessel as a slurry stream, mixed with a liquid comprising water. If the extraction vessel is open-type, the ligno-cellulosic biomass stream is preferably introduced into the extraction vessel by gravity through the free surface of the extraction liquid, for instance by means of a conveyor belt.

A small amount of the carbohydrates of the ligno-cellulosic component may be also solubilized to soluble sugars in the extraction solution, depending on the temperature of the extraction solution and the residence time of the ligno-cellulosic biomass in the extraction vessel.

Preferably, the process conditions are such that the most portion of the water soluble compounds are solubilized in the extraction water while no significant solubilization of the carbohydrates occurs.

Mechanical agitation may be provided for instance by means of paddle wheels to further improve solubilization of the water soluble compounds and to separate the water insoluble components from ligno-cellulosic biomass. The solubilization of the non ligno-cellulosic water soluble compounds is enhanced at high temperature of the extraction solution. The temperature of the extraction solution may be between 30°C and 100°C, preferably between 40°C and 99°C, more preferably between 40°C and 90°C, and most preferably between 50°C and 85°C. Preferably, the temperature is selected to not solubilize the insoluble carbohydrates of the ligno-cellulosic biomass. A fluid at a temperature higher than the extraction solution temperature circulating in a piping system in thermal contact with the extraction solution may be used to heat the extraction solution at the desired temperature. The residence time of the ligno-cellulosic biomass in the extraction solution may be between 30 seconds and 300 minutes, preferably between 1 minute and 20 minutes, more preferably between 2 minutes and 20 minutes, even more preferably between 2 minutes and 15 minutes, and most preferably between 3 and 10 minutes. The residence time may be evaluated by tracing a portion of the ligno-cellulosic biomass in the extraction vessel.

The solubilization step is preferably conducted in an excess of extraction solution with respect to the amount of ligno-cellulosic biomass present in the extraction vessel. Preferably the ratio by weight of the ligno-cellulosic biomass present in the extraction vessel to the extraction liquid in the extraction vessel is less than a value selected from the group consisting of 1 : 1000, 1 :800, 1 :600, 1 :400, 1 :200, 1 : 100, 1 : 70, 1 :50, 1 :30, 1 :20, and 1 : 10. The amount of extraction solution in the extraction vessel is controlled by regulating the flows of streams entering and exiting the extraction vessel.

The extraction solution comprises a total amount of water soluble species which have been accumulated while running the disclosed continuous process. The percent amount of the water soluble species in the extraction solution may be a value in a range selected from the group consisting of 0.25% to 13%, preferably of 0.5% to 8%, more preferably of 0.8% to 5%, and most preferably of 1.25% to 4%.

As the water soluble species comprise charged species and neutral species, the extraction solution may be characterized by the ratio by weight of the total amount of charged species to the total amount of neutral species, which may be in a range selected from the group consisting of 20: 100 to 90: 100, preferably of 40: 100 to 80: 100, and most preferably of 50: 100 to 80: 100.

The pre-soaked ligno-cellulosic biomass stream may be conveyed toward an outlet region of the extraction vessel by means of a mechanical system, which may comprise a paddle conveyor belt, or a paddle wheel, or both.

The pre-soaked ligno-cellulosic biomass is removed from the outlet region of the extraction vessel in the form of a diluted slurry stream with a portion of the extraction solution, and it is drained to separate at least a portion of the free liquid of the pre-soaked ligno-cellulosic biomass slurry stream. Separation occurs preferably under the action of gravity and the drained liquid stream, which is approximately at the same temperature of the extraction solution in the extraction vessel, may be introduced into the extraction vessel. Stated in another way, preferably there is a continuous draining of the liquid stream into the extraction vessel.

A preferred way to remove the pre-soaked ligno-cellulosic biomass stream from the extraction vessel is by means of a mechanical removal system connected to the outlet of extraction vessel and extending to an upper position of the extraction vessel with respect to the gravity.

Preferably, the mechanical removal system comprises a conveyor belt, more preferably a paddle conveyor system, which extracts the pre- soaked ligno-cellulosic biomass slurry stream from an outlet zone of the extraction vessel and drains the dirty liquid stream while lifting the light stream to the upper position. Holes may be suitable located on the conveyor belt to promote draining of the free liquid.

As draining removes most the free liquid in the pre-soaked ligno- cellulosic biomass slurry stream, the pre-soaked ligno-cellulosic biomass stream after draining has a low content of free liquid, which is preferably less than 20%, more preferably less 10%, and most preferably less than 5% weight of the pre-soaked ligno-cellulosic biomass stream after draining on wet basis. In a preferred embodiment, the pre-soaked ligno-cellulosic biomass stream after draining is substantially void of free liquid, that is the free liquid is less than 1% by weight. Free liquid is the liquid which is separated by decanting an aliquot of the light stream after draining in a decanter for 1 hour.

In a preferred embodiment the pre-soaked ligno-cellulosic biomass stream is rinsed with a rinse solution stream, preferably while draining the pre-soaked ligno-cellulosic biomass stream. The rinse solution stream comprises water and it is in general more clean than the extraction solution. In this way, at least a portion of water soluble species which have been solubilized but may adhere to the ligno-cellulosic component are removed from the pre-soaked ligno-cellulosic biomass stream. In this case, the dirty liquid stream comprises the drained rinse solution and has a concentration of water soluble species which is less than the concentration of water soluble species in the extraction solution. The dirty liquid stream may be introduced into the extraction vessel as a dilution stream.

In a preferred embodiment, the pre-soaked ligno-cellulosic biomass stream is rinsed in a limited flow of rinsed solution stream, thereby the disclosed process minimizes the total amount of water needed for treating the ligno-cellulosic biomass. The ratio of the flow of the light stream in Kg/ hour on a dry basis to the flow of the rinse solution stream in Kg/hour is less than 1 :20, preferably less than 1 : 15, more preferably less than 1 : 10, even more preferably less than 1 :7, even yet more preferably less than 1 :5, even yet more preferably less than 1 :3, and most preferably less than 1 : 1.

Preferably the rinse solution stream is injected in a counter-flow configuration with respect to the pre-soaked ligno-cellulosic biomass stream, and it may be injected through one or more injection points while the pre-soaked ligno-cellulosic biomass stream is conveyed by the mechanical removal system.

The rinse solution stream may be at a temperature between 30°C and 100°C, preferably between 40°C and 99°C, more preferably between 40°C and 90°C, and most preferably between 50°C and 85°C.

In an embodiment, the temperature of the rinse solution stream is greater than or equal to the temperature of the extraction liquid and the dirty liquid stream may be introduced into the extraction vessel so as to heat the extraction solution. The pre-soaked ligno-cellulosic biomass is rinsed for a rinsing time which is a value in a range selected from the group consisting of 30 seconds to 300 minutes, 1 minute to 20 minutes, 2 minutes to 20 minutes, 2 minutes to 15 minutes, and 3 minutes to 10 minutes.

The rinsing time may be a significant portion of the residence time of the ligno-cellulosic biomass in the extraction vessel. Preferably, the rinse time is in a range between 1% and 80%, more preferably between 5% and 70%, even more preferably between 10% and 60%, and most preferably between 20% and 50% of the residence time.

Thereby, an additional portion of the non-ligno-cellulosic water soluble compounds may be further solubilized during rinsing and removed from the pre-soaked ligno-cellulosic biomass.

Even if the pre-soaked ligno-cellulosic biomass stream contains few or no free liquid, the moisture content is still high, being the ligno- cellulosic biomass soaked with the extraction solution. The moisture content may be in the range 70% to 95%, preferably of 70% to 90%, more preferably of 75% to 90%, and most preferably of 75% to 85%.

An additional portion of soaking liquid may be removed from the pre- soaked ligno-cellulosic biomass stream preferably by means of a continuous compression device. A preferred device is a compression screw, located in a cylindrical housing having an annular filter screen to remove liquids while continuously conveying the clean ligno-cellulosic biomass stream. After the continuous pressing, the moisture content of the pre-soaked ligno-cellulosic biomass stream is in the range of 40% to 75%, preferably of 40% to 70%, more preferably of 45% to 65%, and most preferably of 45% to 60%.

The pre-soaked ligno-cellulosic biomass is then passed to the second hydrothermal treatment.

Preferred second hydrothermal treatment of the pre-soaked ligno- cellulosic feedstock The liquid biomass feed stream is derived from the pre-soaked ligno- cellulosic feedstock by means of a treatment, or pre-treatment, of the pre-soaked ligno-cellulosic feedstock.

The pre-treatment of the pre-soaked ligno-cellulosic biomass is used to solubilize and remove carbohydrates, mainly xylans and glucans, from the pre-soaked ligno-cellulosic feedstock, and at the same time the concentrations of harmful inhibitory by-products such as acetic acid, furfural and hydroxymethyl furfural remain substantially low.

Pre-treatment techniques which may be used are well known in the art and include physical, chemical, and biological pre-treatment, or any combination thereof. In preferred embodiments the pre-treatment of pre-soaked ligno-cellulosic biomass is carried out as a batch or continuous process.

Physical pre-treatment techniques include various types of milling/ comminution (reduction of particle size), irradiation

Comminution includes dry, wet and vibratory ball milling.

Although not needed or preferred, chemical pre-treatment techniques include acid, dilute acid, base, organic solvent, lime, ammonia, sulfur dioxide, carbon dioxide, pH-controlled hydro thermolysis, wet oxidation and solvent treatment.

If the chemical treatment process is an acid treatment process, it is more preferably, a continuous dilute or mild acid treatment, such as treatment with sulfuric acid, or another organic acid, such as acetic acid, citric acid, tartaric acid, succinic acid, or any mixture thereof. Other acids may also be used. Mild acid treatment means at least in the context of the invention that the treatment pH lies in the range from 1 to 5, preferably 1 to 3.

In a specific embodiment the acid concentration is in the range from 0.1 to 2.0 % wt acid, preferably sulfuric acid. The acid is mixed or contacted with the pre-soaked ligno-cellulosic biomass and the mixture is held at a temperature in the range of around 160-220 °C for a period ranging from minutes to seconds. Specifically the pre-treatment conditions may be the following: 165- 183 °C, 3- 12 minutes, 0.5- 1.4% (w/w) acid concentration, 15-25, preferably around 20% (w/w) total solids concentration. Other contemplated methods are described in U.S. Pat. Nos. 4,880,473, 5,366,558, 5, 188,673, 5,705,369 and 6,228, 177.

Wet oxidation techniques involve the use of oxidizing agents, such as sulfite based oxidizing agents and the like. Examples of solvent treatments include treatment with DMSO (Dimethyl Sulfoxide) and the like. Chemical treatment processes are generally carried out for about 5 to about 10 minutes, but may be carried out for shorter or longer periods of time.

In an embodiment both chemical and physical pre-treatment is carried out including, for example, both mild acid treatment and high temperature and pressure treatment. The chemical and physical treatment may be carried out sequentially or simultaneously.

The current strategies of thermal treatment are subjecting the pre- soaked ligno-cellulosic material to temperatures between 1 10-250°C for 1-60 min e.g.:

Hot water extraction

Multistage dilute acid hydrolysis, which removes dissolved material before inhibitory substances are formed

Dilute acid hydrolysis at relatively low severity conditions Alkaline wet oxidation

Steam explosion

Almost any pre-treatment with subsequent detoxification. If a hydrothermal pre-treatment is chosen, the following conditions are preferred:

Pre-treatment temperature: 1 10-250°C, preferably 120-240°C, more preferably 130-230°C, more preferably 140-220°C, more preferably 150-210°C, more preferably 160- 200°C, even more preferably 170-200°C or most preferably 180-200°C.

Pre-treatment time: 1-60 min, preferably 2-55 min, more preferably 3- 50 min, more preferably 4-45 min, more preferably 5-40 min, more preferably 5-35 min, more preferably 5-30 min, more preferably 5-25 min, more preferably 5-20 min and most preferably 5- 15 min.

Dry matter content after pre-treatment is preferably at least 20% (w/w). Other preferable higher limits are contemplated as the amount of biomass to water in the pre-treated ligno-cellulosic feedstock be in the ratio ranges of 1 :4 to 9: 1 ; 1 :3.9 to 9: 1 , 1 :3.5 to 9: 1 , 1 :3.25 to 9: 1 , 1 :3 to 9: 1 , 1 :2.9 to 9: 1 , 1 :2 to 9: 1 , 1.15 to 9: 1 , 1 : 1 to 9: 1 , and 1 :0.9 to 9: 1.

A preferred pretreatment of the pre-soaked ligno-cellulosic biomass include a soaking of the pre-soaked ligno-cellulosic biomass feedstock and optionally a steam explosion of at least a part of the soaked ligno-cellulosic biomass feedstock. The soaking occurs in a substance such as water in either vapor form, steam, or liquid form or liquid and steam together, to produce a product. The product is a soaked biomass containing a soaking liquid, with the soaking liquid usually being water in its liquid or vapor form or some mixture. This soaking can be done by any number of techniques that expose a substance to water, which could be steam or liquid or mixture of steam and water, or, more in general, to water at high temperature and high pressure. The temperature should be in one of the following ranges: 145 to 165°C, 120 to 210°C, 140 to 210°C, 150 to 200°C, 155 to 185°C, 160 to 180°C. Although the time could be lengthy, such as up to but less than 24 hours, or less than 16 hours, or less than 12 hours, or less than 9 hours or less than 6 hours; the time of exposure is preferably quite short, ranging from 1 minute to 6 hours, from 1 minute to 4 hours, from 1 minute to 3 hours, from 1 minute to 2.5 hours, more preferably 5 minutes to 1.5 hours, 5 minutes to 1 hour, 15 minutes to 1 hour.

If steam is used, it is preferably saturated, but could be superheated. The soaking step can be batch or continuous, with or without stirring, and it may also include the addition of other compounds, e.g. H2SO4, NH3, in order to achieve higher performance later on in the process. The product comprising the soaking liquid, or soaked liquid, is then passed to a separation step where at least a portion of the soaking liquid is separated from the soaked biomass. The liquid will not completely separate so that at least a portion of the soaking liquid is separated, with preferably as much soaking liquid as possible in an economic time frame. The liquid from this separation step is known as the soaked liquid stream comprising the soaking liquid. The soaked liquid will be the liquid used in the soaking, generally water and the soluble species of the feedstock. These water soluble species comprise glucan, xylan, galactan, arabinan, and their monomers and oligomers. The solid biomass is called the first solid stream as it contains most, if not all, of the solids.

The separation of the soaked liquid can again be done by known techniques and likely some which have not yet been invented. A preferred piece of equipment is a press, as a press will generate a liquid under high pressure.

The first solid stream may then optionally be steam exploded to create a steam exploded stream, comprising solids. Steam explosion is a well- known technique in the biomass field and any of the systems available today and in the future are believed suitable for this step. The severity of the steam explosion is known in the literature as Ro, and is a function of time and temperature and is expressed as Ro = texp[(T- 100)/ 14.75] with temperature, T expressed in Celsius and time, t, expressed in minutes.

The formula is also expressed as Log(Ro), namely Log(Ro) = Ln(t) + [(T- 100)/ 14.75] .

Log(Ro) is preferably in the ranges of 2.8 to 5.3, 3 to 5.3, 3 to 5.0 and 3 to 4.3.

The steam exploded stream may be optionally washed at least with water and there may be other additives used as well. It is conceivable that another liquid may be used in the future, so water is not believed to be absolutely essential. At this point, water is the preferred liquid. The liquid effluent from the optional wash may be added to the soaked liquid stream. This wash step is not considered essential and is optional. The washed exploded stream is then processed to remove at least a portion of the liquid in the washed exploded material. This separation step is also optional. The term at least a portion is removed, is to remind one that while removal of as much liquid as possible is desirable (preferably by pressing), it is unlikely that 100% removal is possible. In any event, 100% removal of the water is not desirable since water is needed for the subsequent hydrolysis reaction. The preferred process for this step is again a press, but other known techniques and those not invented yet are believed to be suitable. The liquid products separated from this process may be added to the soaked liquid stream. The liquid biomass feed stream comprises the soaking liquid and may further comprise the additional liquid streams produced in the treatment.

EXPERIMENTAL 1. Preparation of the liquid biomass feed streams

Wheat straw, pre-treated at the indicated conditions, is the reference ligno-cellulosic feedstock used for validating the invention. Other ligno- cellulosic feedstocks pre-treated at different conditions were also tested. 1.1. First Hydrothermal treatment

Wheat straw was subjected to two different pre-soaking hydrothermal treatments to solubilize a portion of the non ligno-cellulosic water soluble compounds.

In the first pre-soaking treatment (sample WEI), wheat straw was continuously introduced into a closed vessel having a parallelepiped shape, at an inlet position located close to one end of vessel. The feedstock was introduced at a rate of 40 Kg/h on a dry basis, and it was transported to the outlet of the vessel by means of a conveyor belt.

The feedstock was treated in at 80°C for a residence time of 3 minutes in an extraction solution comprising water and water soluble species accumulated during the continuous treatment.

The volume of the extraction solution was 2 m ³ , and fresh water was continuously added to the vessel at a feedstock : liquid ratio by weight on a dry basis of 1 :3. The soaked feedstock was removed from the outlet of the vessel by means of an inclined conveyor belt and drained under the effect of gravity. The final moisture content (or dry matter) was 80%.

The second pre-soaking treatment (sample WE2) was conducted as the first pre-soaking treatment, but it was counter-washed with water while it was drained on the inclined conveyor belt. Counter-washing water was spread on the feedstock through 6 nozzles distributed along the conveyor belt length, at a temperature of 80°C. Residence time on the conveyor belt was 2 minutes. The feedstock was then compressed by means of a compression screw to further remove pre-soaking liquid. The final moisture content (or dry matter) was 60%. The composition on dry basis of the pre-soaked biomasses and of the raw material, which was not pre-soaked, is reported in table 1. The table shows that the pre-soaking treatments removed a relevant portion of the water soluble components, and that the second presoaking treatment, characterized by a counter-washing and squeezing steps, was more effective than the first soaking treatment. It is noted that the percent amount of the water insoluble components increases due to the removal of the water soluble components.

Table 1. Composition of CE-raw, WEl- lHT and WE2- 1HT biomasses as entering the second hydro thermal treatment.

1.2. Second hydrothermal treatment

The pre-soaked wheat straw was introduced into a continuous reactor and subjected to a soaking treatment at a temperature of 158°C for 65 minutes. The soaked mixture was separated in a soaking liquid and a fraction containing the solid treated raw material by means of a press.

The soaking liquid was then subjected to concentration by means of vacuum evaporation.

The dry matter of the soaking liquid after concentration was about 12% and the pH was 4.5. The soaking liquids are the liquid biomass feed streams used for validating the disclosed process.

For control, a portion of feedstock was subjected to the same hydrothermal treatment without any preceding presoaking treatment (sample CE).

Table 2. Composition of the three liquid biomass feed streams Table 2 reports the detailed composition of the soaking liquids, of the samples CE-2HT, WE1-2HT and WE2-2HT, which are the liquid biomass feed streams of the disclosed process. The data show that the percent amount of total sugars, increases significantly in the samples subjected to pre-soaking, and the increase is especially relevant in the case of the presoaking with counter- washing and squeezing steps. In particular, xylooligomers solubilization is increased. Thereby, the presoaking increases the solubilization of insoluble sugar polymers, in particular the solubilization of xylans. The data also show that the percent amount of water soluble compounds different from soluble sugars is strongly reduced by means of the presoaking treatment, and the effect is more evident in the case of the counter-washing and squeezing steps.

In table 3, the concentration of the main water soluble compounds different from soluble sugars relative to the total concentration of soluble sugars on dry basis is reported, to emphasize the combined effect of the removal of water soluble compounds in the first hydrothermal treatment and the increase of sugar solubilization in the second hydrothermal treatment.

Table 3. Concentration of the main water soluble compounds different from soluble sugars relative to the total concentration of soluble sugars on dry basis

2. Hydrolysis of the liquid biomass feed streams

2.1. Hydrolysis by decationization 2.1.1. Decationization of the liquid biomass feed streams

A portion of the three liquid biomass feed streams were subjected to decationization step according to the following procedure.

The decationization step was performed in a decationization column having a working volume of 60 liters. The column was filled with a strong cationic resin Relite EXC14 supplied by Mitsubishi Chemical Corporation. The resin was subjected to an activation step by means of HC1. Briefly, 60 gr of HC1 per liter of resin is used in form of a solution at 5% concentration, at a temperature of 25°C and at a rate of 3 BV/h. The resin was then washed with abundant water. After resin activation, the liquid biomass feed streams were inserted at a rate of 4 BV/h at 25°C.

The decationized liquid biomass feed streams extracted from the decationization column are the acidic liquid biomass feed stream of the disclosed process. The pH of the decationized liquid biomass feed streams were 0.87.

2.1.2. Hydrolysis of the acidic streams

The acidic liquid biomass feed streams were hydrolyzed in a 3 liter PFR reactor according to the following procedure. Residence time and temperature of each test were determined taking into account the different pH of the streams. The hydrolysis conditions are reported in Table 4. PH Temperature (°C) Time (min)

CE-dec 1.01 157 3

WEl-dec 1.40 157 6

WE2-dec 1.78 157 15

Table 4. Parameters oh hydrolysis conditions of the acidic liquid biomass feed streams.

After hydrolysis, the liquid sugar streams were concentrated by vacuum evaporation to reach a dry matter content by weight of about 37%. The liquid sugar streams hydrolyzed by decationization are indicated as CE- dec, WEl-dec and WE2-dec.

2.2. Enzymatic hydrolysis

A portion of the liquid biomass feed streams CE-2HT and WE2-2HT were subjected to enzymatic hydrolysis. The liquid biomass feed streams were inserted into a bioreactor, agitated by means of an impeller and heated till reaching a temperature of 50°C. pH was corrected to 5 by means of a KOH solution.

Enzymatic hydrolysis was conducted by inserting an enzymatic cocktail rich in hemicellulose (Cellic HTEC3 by Novozymes) at a determined concentration of 2mg of protein per gram of xylo-oligomers contained in the liquid biomass feed streams.

Enzymatic hydrolysis was conducted for 48hours.

The enzymatically hydrolyzed streams were subjected to a softening step conducted with the same procedure of decationization step, with the difference that in this case the strong cationic resin was used in Na form.

After softening, the liquid sugar streams were concentrated by vacuum evaporation to reach a dry matter content by weight of about 37%. The liquid sugar streams hydrolyzed by enzymatic hydrolysis are indicated as CE-enz and WE2-enz.

3. Neutralization of the liquid sugar streams

The liquid sugar streams were cooled at 25°C. Sodium hydroxide was added and mixed to the liquid sugar streams in a sufficient amount to raise the pH to 6.0.

The composition of all the liquid sugar streams after neutralization are reported in Table 5.

The compositional analyses have been performed to identify a long list of compounds, in order to highlight the effect of presoaking step.

The compounds have been grouped into general categories to simplify the reading of the tables and showing the categories of compounds relevant for the purification process. Sugar compounds are grouped into monomeric sugars (glucose, xylose, arabinose, fructose, mannose, galactose, ribose) and oligomeric sugars (cellobiose, xylobiose, sucrose, gluco-oligomers, xylo-oligomers, arabino-oligomers). It is noted that dimers are considered as oligomeric sugars in the framework of the present disclosure.

The category "other neutral compounds" includes glycerol, 5-HMF and furfural, which are neutral compounds different from monomeric and oligomeric sugars, which are neutral. These compounds are usually considered important for conversion process of the monomeric sugars to other compounds, following the disclosed purification process, as they may poison the catalyst and reduce efficiency. It is noted that other neutral compounds may be present also in other categories.

The category "organic acids" includes formic acid, lactic acid, acetic acid, levulinic acid and other carboxylic acid. These acids have been identified by measuring the corresponding anions. The other carboxylic acids include Pyruvic acid, D-Galacturonic acid and Glucuronic acid. The category "cations" include monovalent inorganic cations (sodium, ammonium and potassium) and non-monovalent inorganic cations (magnesium and calcium). These cations are considered for calculating the two relevant ratios of the disclosed process, namely the ratio of the total ionic equivalents of cations different from Sodium to the total ionic equivalents of cations and the ratio of the total ionic equivalents of non- monovalent cations to the total ionic equivalents of cations in the liquid sugar stream.

Table 5. Composition of the liquid sugar streams after neutralization, grouped into different categories

The category "other anions" indicates the following anions not comprised in the "organic acids" category: chloride, nitrite, bromide, nitrate, sulfate and phosphate.

In the table, acetyls, organic nitrogen and phenolic compounds are also reported. Phenolic compounds, which are also considered detrimental for catalyst poisoning, include a long list of compounds, namely: Catechin, 4-hydroxybenzaldehyde, vanillic acid, Caffeic acid, Syringic acid, Vanillin, Syringaldehyde, Acetovanillone, p-coumaric acid, Acetosyringone and ferulic acid.

The compounds not present in the previous categories are grouped in the generic category "other compounds" and are calculated as the complementary to 100% on the measured compounds.

In the case of CE-enz and WE2-enz, other carboxylic acids were not quantified separately and thereby are inserted into the "other compounds" category.

Compositions are given as percent amount on a dry basis.

It is noted that, in all the streams, the ratio of the total ionic equivalents of cations different from Sodium to the total ionic equivalents of cations is less than 3%. In figure 1 the amount of base used to neutralized the liquid sugar streams by decationization, relative to the amount of sugar, is reported.

It is noted that the amount of base used in WEI -dec is 65% of the base of CE-dec, while the amount of base added in WE2-dec is only 35% of the base added in CE-dec. The disclosed process thereby strongly reduces the costs associated to neutralization.

4. Chromatographic separation of the liquid sugar streams

The liquid sugar streams were subjected to batch chromatographic separation to produce the purified liquid sugars streams of the disclosed process.

A first portion of the liquid sugar streams was subjected to a batch chromatographic separation in a 60 liter column filled with UBK530 by Mitsubishi Chemical Corporation, provided in Sodium form, according to the following procedure: 5% of BV (31iter) of the liquid sugar stream at a temperature of 50°C was inserted into the column at a flow rate of 0.5BV/h, followed by inserting of elution water at the same flow rate for a total bed volume of IBV. Different fractions were collected at step of 0.05BV. The resin has a ion exclusion effect, as evidenced by electrical conductivity measurements, and also a size exclusion effect.

In figure 2 and figure 3 it is represented the percent dry matter present in the relevant chromatographic fractions for the decationization and enzymatic hydrolysis samples, respectively. The sum of all the percent dry matter values is the total dry matter of the feeding liquid sugar streams. It is noted that the chromatograms present two evident peaks, wherein the left peak is due mainly to non-monosaccharide compounds and the right peak is due mainly to monosaccharide compounds. In both the figures, the first hydrothermal treatment, which removes a relevant portion of non ligno-cellulosic water soluble compounds, greatly reduces the area of the first peak and the two peaks are more defined and separated with respect to the control experiments.

In table 6 the compositions of the chromatographic fractions on a dry basis, the dry matter content, the purity ratio and the recovery ratio are reported for the samples produced by decationization.

The first seven fractions are omitted as they contain substantially only water.

In table 7 the compositions of the chromatographic fractions on a dry basis, the dry matter content, the purity ratio and the recovery ratio are reported for the samples produced by enzymatic hydrolysis.

The different combinations of the chromatographic fractions may be considered as the purified liquid sugar stream, being the other fractions considered as residual streams.

In the case of CE-dec, for having a high purity ratio, it is possible to combine chromatographic fractions from 13 to 15, but the recovery ratio is 87.5% (purity ratio:90.0%).

The purity ratio of the combined chromatographic fractions is calculated as the average of the purity of the single fractions weighted on the dry matter of the fractions, taking into account the mass density. For having a high recovery ratio, for instance greater than 90%, it is possible to combine fractions 12- 15 (recovery ratio: 92.0%, purity ratio: 81.8%), fractions 13- 16 (recovery ratio: 92.0%, purity ratio: 88.1%) or fractions 12- 16 (recovery ratio: 96.5%, purity ratio: 80.6%), thereby reducing the purity of the recovered combined stream.

In the case of WEI -dec, it is possible to combine chromatographic fractions from 13 to 18, to have a recovery ratio of 92.4% and a purity ratio of 90.0%. For having a higher recovery ratio, it is possible to combine fractions 12- 18, to obtain a recovery ratio of 99.7%, and a still high purity ratio of 87.9% .

In the case of WE2-dec, it is possible to combine chromatographic fractions from 13 to 18, to have a purity ratio of 92.3% and a recovery ratio of 95.6%. For having a higher recovery ratio, it is possible to combine fractions 12- 18, to obtain an impressive recovery ratio of 99.0%, and a still high purity ratio of 88.2%.

In the case of enzymatic hydrolysis, considering CE-enz, by combining fractions 13- 16 it is possible to obtain a recovery ratio of 80.8%, which is low, and a high purity ratio of 91.2%, instead by combining fractions 12- 16 it is possible to obtain a high recovery ratio of 95.4%, but a low purity ratio of 71.3% .

Considering WE2-enz, by combining fractions 13- 16 it is possible to obtain a high recovery ratio of 93.8% and a high purity ratio of 92.3%, instead by combining fractions 12- 16 it is possible to obtain a high recovery ratio of 98.5%, but a still high purity ratio of 86.9%. It is pointed out that the surprising and outstanding results of the disclosed process are represented by the high purity of the purified liquid sugar stream obtained with a high recovery ratio, starting from a liquid biomass feedstream which contains a large amount of impurities. The presoaking reduces the amount of impurities in the liquid sugar stream and greatly improves the chromatographic separation. According to inventors knowledge, no other process disclosed in the prior art reaches this result. Without being limited by any interpretation, it is believed that this result is due to the combination of the presoaking step and the specific chromatographic separation of the disclosed process.

After neutralization by NaOH addition, very few cations different from Sodium are present in the liquid sugar streams.

These feature greatly improves the lifetime, or durability, of the Na- based separation resin. Thereby, the disclosed process is effective in producing a high purity liquid sugar stream without wasting significant amount of sugars, as needed to improve the economics in industrial production of biochemical products, such as ethylene glycol.

liquid sugar

stream 4.21% 33.05% 9.34% 53.39% 37.09% 46.61% 100.00%

8 0.00% 0.00% 0.00% 100.00% 2.22% 0.00% 0.00%

9 0.00% 0.00% 0.00% 100.00% 3.24% 0.00% 0.00%

10 0.00% 0.00% 0.00% 100.00% 5.20% 0.00% 0.00%

11 1.78% 0.00% 0.00% 98.22% 3.44% 1.78% 0.30%

12 23.08% 40.41% 4.75% 31.76% 2.11% 68.24% 7.26%

Fraction 13 10.02% 69.66% 10.86% 9.46% 7.58% 90.54% 34.57% number 14 4.79% 66.39% 19.29% 9.53% 7.69% 90.47% 34.33%

15 2.56% 58.65% 27.90% 10.89% 3.76% 89.11% 16.31%

16 2.36% 54.71% 35.45% 7.47% 1.27% 92.53% 5.69%

17 0.00% 42.47% 36.56% 20.98% 0.32% 79.02% 1.21%

18 0.00% 24.52% 25.58% 49.89% 0.13% 50.11% 0.31%

19 0.00% 0.00% 7.69% 92.31% 0.08% 7.69% 0.03%

Other

monomelic Other Purity ratio Recovery

WE2-dec Glucose Xylose sugars compounds DM (%) (%) ratio (%) liquid sugar

stream 3.41% 39.46% 10.19% 46.94% 37.09% 53.06% 100.00%

8 0.00% 0.00% 0.00% 100.00% 2.36% 0.00% 0.00%

9 0.39% 0.00% 0.00% 99.61% 4.39% 0.39% 0.07%

10 0.54% 0.00% 0.00% 99.46% 6.28% 0.54% 0.15%

11 1.04% 0.00% 0.00% 98.96% 4.00% 1.04% 0.18%

12 11.04% 23.94% 4.72% 60.30% 2.04% 39.70% 3.44%

Fraction 13 8.09% 73.38% 9.55% 8.97% 5.44% 91.03% 21.35% number 14 5.39% 70.93% 16.47% 7.20% 6.98% 92.80% 28.33%

15 3.75% 70.82% 20.87% 4.55% 5.64% 95.45% 23.27%

16 2.92% 65.85% 22.80% 8.43% 3.57% 91.57% 13.98%

17 1.95% 59.44% 25.58% 13.03% 1.73% 86.97% 6.38%

18 0.00% 56.66% 29.69% 13.65% 0.62% 86.35% 2.26%

19 0.00% 40.84% 27.56% 31.60% 0.21% 68.40% 0.60%

Table 6. Composition of the fractions obtained by batch chromatographic separation of liquid sugar streams CE-dec, WEI -dec and WE2-dec.

number 9 1.09% 1.15% 0.00% 97.75% 5.95% 2.25% 1.04%

10 0.71% 1.11% 0.00% 98.18% 9.15% 1.82% 1.31%

11 1.01% 1.65% 0.00% 97.34% 9.29% 2.66% 1.96%

12 8.16% 24.22% 0.00% 67.61% 5.84% 32.39% 14.68%

13 10.83% 80.57% 2.94% 5.65% 6.05% 94.35% 44.38%

14 4.85% 79.83% 9.57% 5.75% 4.54% 94.25% 33.04%

15 0.93% 45.72% 23.24% 30.12% 0.62% 69.88% 3.26%

16 0.00% 4.59% 0.00% 95.41% 0.23% 4.59% 0.08%

17 0.00% 5.54% 0.00% 94.46% 0.14% 5.54% 0.06%

18 0.00% 6.69% 0.00% 93.31% 0.10% 6.69% 0.05%

Other

monomeric Other Purity ratio Recovery

WE2-enz Glucose Xylose sugars compounds DM (%) (%) ratio (%) liquid sugar

stream 3.46% 41.67% 4.61% 50.26% 37.09% 49.74% 100.00%

8 0.58% 0.62% 0.00% 98.80% 1.83% 1.20% 0.10%

9 0.57% 0.55% 0.00% 98.88% 5.39% 1.12% 0.28%

10 0.53% 0.53% 0.00% 98.94% 7.07% 1.06% 0.35%

11 1.15% 0.98% 0.00% 97.87% 4.85% 2.13% 0.48%

12 18.19% 41.50% 0.21% 40.10% 1.74% 59.90% 4.75%

Fraction 13 8.37% 80.56% 2.00% 9.06% 7.42% 90.94% 31.67% number 14 3.84% 82.37% 7.88% 5.91% 9.40% 94.09% 41.96%

15 1.24% 69.80% 21.06% 7.90% 4.53% 92.10% 19.32%

16 0.87% 24.16% 46.14% 28.83% 0.25% 71.17% 0.80%

17 0.00% 6.47% 0.00% 93.53% 0.19% 6.47% 0.06%

18 0.00% 10.54% 8.80% 80.67% 0.11% 19.33% 0.09%

19 0.00% 14.37% 7.94% 77.69% 0.13% 22.31% 0.13%

Table 7. Composition of the fractions obtained by batch chromatographic separation of liquid sugar streams CE-enz and WE2- enz.