Continuous process for fractionating a suspension

Abstract

The invention relates to a continuous process for fractionating a suspension chosen from a microalgal biomass or milk.

Claims

1. A process of fractionation of a suspension chosen from a microalgal biomass or milk, comprising the following steps: a) contacting the suspension with a solvent consisting in supercritical CO.sub.2 in a contactor, the suspension and the supercritical CO.sub.2 being introduced in a counter current mode, b) continuously recovering a supercritical CO.sub.2 rich phase exiting from the top of the contactor, after its contacting with the suspension, the supercritical CO.sub.2 rich phase containing lipids extracted from the suspension, and a raffinate exiting from the bottom of the contactor, c) continuously separating the neutral lipids and/or the compounds of interest contained in the supercritical CO.sub.2 rich phase, and the raffinate, d) the process comprising no step of disrupting the cells of the algal biomass before step a) is carried out.

2. The process of claim 1, wherein the suspension is a microalgal biomass, said microalgal biomass having a concentration of microalgae comprised between 0.5 and 200 g.Math.L.sup.1 of microalgae.

3. The process of claim 1, wherein the suspension is milk, said milk being from animal or vegetal origin.

4. The process of claim 1, wherein the contactor is a fractionation column, in particular a packed fractionation column, or a settler-mixer.

5. The process of claim 4, wherein the packing of the fractionation column is chosen among random packings and structured packings.

6. The process of claim 1, wherein the fractionation is carried out at a pressure comprised between 8 and 30 MPa and at a temperature comprised between 35 and 70 C.

7. The process of claim 1, carried out with the massic ratio [flow rate of supercritical CO.sub.2]/[flow rate of suspension] comprised between 2/1 and 250/1.

8. The process of claim 1, carried out with: the flow rate of supercritical CO.sub.2 comprised between 3 and 20 kg h.sup.1 and the flow rate of suspension comprised between 0.08 and 1.5 kg h.sup.1; the flow rate of supercritical CO.sub.2 of about 50 kg h.sup.1 and the flow rate of suspension of about 5 kg h.sup.1; or the flow rate of supercritical CO.sub.2 of about 600 kg h.sup.1 and the flow rate of suspension of about 50 kg h.sup.1.

9. The process of claim 1, where the separated supercritical CO.sub.2 is recirculated in the fractionation column.

10. The process of claim 1, wherein the neutral lipids are chosen from triglycerides, diglycerides, monoglycerides, free fatty acids and their mixtures, the neutral lipids being in particular -3 fatty acids.

11. The process of claim 1, wherein the compounds of interest antioxidants, carbohydrates or pigments.

12. The process of claim 1, wherein the contacting step a) is performed in presence of a polar modifier and/or an esterification agent.

Description

FIGURES

(1) FIG. 1 is a schematic diagram of a laboratory-scale fractionation unit that may be used in the process of the invention. This fractionation unit includes a packed column noted 1, with an internal diameter of 19 mm and 2 m height with a viewing cell (not represented) located at the bottom of column 1, below the solvent (supercritical CO.sub.2) injection nozzle (not represented). The temperature of column 1 is controlled by means of two independent heating jackets, noted 2 and 3.

(2) C1: cooler; H1, H2: heaters; P1, P2: High-pressure pumps, S1: separator; PV: back-pressure regulator; V1-V7: valves(1) column; (2), (3): column sections.

(3) FIG. 2 is a schematic diagram of a pilot or industrial-scale fractionation unit that may be used in the process of the invention. This fractionation unit includes a packed column noted 1. The temperature of column 1 is controlled by means of four independent heating jackets, noted 2-5.

(4) C1, C2: coolers; H1-H4: heaters; P1, P2: High-pressure pumps, S1,S2,R2: separators; R1: buffer tank; PV: back-pressure regulator; V1-V12: valves(1) column; (2)-(5) column sections.

EXAMPLE 1: SUPERCRITICAL FRACTIONATION OF NEUTRAL LIPIDS (MAINLY TRIGLYCERIDES AND FATTY ACIDS) FROM A MACROALGAL SUSPENSION

(5) The process of the invention will be described in reference to FIG. 1 annexed which is a schematic diagram of a fractionation unit used in the process of the invention.

(6) This fractionation unit includes a packed column noted 1 in FIG. 1, with an internal diameter of 19 mm and 2 m height with a viewing cell (not represented) located at the bottom of column 1, below the solvent (supercritical CO.sub.2) injection nozzle (not represented).

(7) The column is able to withstand pressures up to 30 MPa, and its temperature (from 35 to 70 C.) is controlled by means of two independent heating jackets, noted 2 and 3 in FIG. 1. The column is filled with 10 mm Interpack random packings from VFF (Germany), with a measured apparent density of 588 kg m.sup.3, a specific surface area of 580 m.sup.1 and a void fraction (equivalent porosity) of 0.917 in order to improve mass transfer efficiency.

(8) During operation, carbon dioxide which is introduced through a valve noted V6 in FIG. 1, under 4.5 MPa is cooled to 278 K (5 C.) in a double tube heat exchanger, noted C1 in FIG. 1, before being pumped and then heated to the working temperature, by a heater noted H2 in FIG. 1. A high-pressure piston pump, noted P2 in FIG. 1, from Separex (France) with a top capacity of 15 Lh.sup.1 of liquid carbon dioxide and a maximum attainable pressure of 35 MPa delivers the supercritical CO.sub.2 at the suitable flow rate. In this example, the working temperature is 60 C. and the working pressure is 25 MPa. Those conditions allow the recovery of neutral lipids with high separation efficiency.

(9) However, the temperature may vary from 35 to 70 C. and the pressure may vary from 10 to 30 MPa.

(10) The CO.sub.2-over-feed ratio in the present example varies between 20 and 150 corresponding to a CO.sub.2 flow rate set at 12 kg h.sup.1 and feed flow rate varying from about 0.08 up to 0.55 kg h.sup.1.

(11) However, this CO.sub.2 flow rate may vary from 5 to 15 kg h.sup.1.

(12) When supercritical CO.sub.2 leaves column 1, the overhead current is depressurized through a backpressure regulator, noted PV in FIG. 1, to recover the extract in a pressurized cyclonic separator, noted S1 in FIG. 1. The extract is heated using the heater H1 placed between the exit of the column and the separator to ensure to maintain the desired temperature.

(13) The CO.sub.2 rich phase is no longer a supercritical fluid, it is then recycled and condensed into the cooler C1 to reduce carbon dioxide consumption.

(14) If the separation between the extracted compounds and CO.sub.2 is not complete, an additional apparatus may be used in order to purify the CO.sub.2 rich phase. For example, activated carbon may be used to ensure the complete gas separation.

(15) In this example, the concentration of the algal biomass is of 100 g L.sup.1.

(16) However, this concentration may vary from 0.5 to 200 g L.sup.1.

(17) The liquid mixture which is fed by a Gilson 307 HPLC piston pump, noted P1 in FIG. 1, which has a 20 mL min.sup.d (about 1.2 kg h.sup.1) maximum capacity, through a value noted V1 in FIG. 1.

(18) The supercritical CO.sub.2 flow rate is controlled by a Rheonik RHE 14 mass flowmeter (Germany) (not represented in FIG. 1), while the algal biomass feed flow rate is directly controlled by the pump P1 speed.

(19) The extract containing neutral lipids and the raffinate are collected from the bottom of the cyclonic separator S1 and of column 1, respectively, by manual regulation of the corresponding valves, noted V2 and V7 in FIG. 1.

(20) The microalgal biomass is introduced in the top of column 1 via a pump noted P1, and a valve noted V1 in FIG. 1.

(21) The supercritical solvent (supercritical CO.sub.2) is introduced at the bottom of the column (1).

(22) Thus, the supercritical CO.sub.2 and the algal biomass flow in the column at counter-current. Supercritical CO.sub.2 solubilizes all or part of the neutral lipids and/or the products with a high added value such as pigments or antioxidants. If pure carbon dioxide is used, the neutral lipids are the most soluble compounds and will be then preferentially be extracted from the algal cells. Those neutral lipids can be tri-, di- or monoglycerides as well as free fatty acids. Even if pure supercritical CO.sub.2 is used, since water can act as a co-solvent, -caroten can also be solubilized in the light phase. If a polar modifier (ethanol, for instance) is added to supercritical CO.sub.2, some more polar compounds will be solubilized in the fluid phase, as astaxanthin or lutein, and the -caroten extraction will be enhanced.

(23) The algal biomass, from which the neutral lipids and the compounds of interest have been extracted, is recovered at the bottom of the column through a valve, noted V7 in FIG. 1.

(24) In the present case, since feed flow rate varies from 0.08 up to 0.55 kg h.sup.1, a quantity of neutral lipids varying from 16 to 100 g can be recovered after one hour of production for a biomass containing 20 wt % (of dry matter) of neutral lipids and for a complete recovery. Those quantities can be as high as 10 kg (of neutral lipids) per hour of production at a larger scale, for example with a 8 m high column and with an internal diameter of 126 mm.

(25) When supercritical CO.sub.2 is depressurized, it returns to its gaseous state so that the oil containing the lipids and any compounds of interest are spontaneously separated.

(26) In the present case, the oil contained 20 wt % (of dry matter) of neutral lipids and up to 2 wt % (of dry matter) of polar compounds (antioxidants).

(27) The glycerides (tri-, di- and mono-), the free fatty acids, antioxidants and the other compounds of interest can be separated from each other.

(28) The triglycerides are separated from the antioxidants using different separators with different conditions of pressure and temperature. As for example, the operating conditions of the first separator may be set at 20 MPa and 40 C. allowing the recovery of pigments. The operating conditions of the second separator may be set at 6 MPa and 40 C. allowing the recovery of glycerides.

EXAMPLE 2: DELIPIDATION OF MILK

(29) Depending on its origin, milk is constituted of about 87% of water and 13% of dry extract. The dry extract represents about 130 g.Math.L.sup.1 and contains on the average of 35-45 g of fat. About 98.5% of fat fraction are simple lipids in suspension: 95-96% are triglycerides, 3% are diglycerides, the rest are monoglycerides.

(30) Since the water content and the composition of neutral lipids are almost the same for milk and microalgae suspension, the apparatus, method and operating conditions presented in example 1 can be applied for both types of feed.

(31) As for example, the working temperature in the column is fixed at 60 C. and the working pressure is 25 MPa.

(32) The CO.sub.2-over-feed mass ratio in the present example (for ex. for a 2 m high column) varies between 20 and 150 corresponding to a CO.sub.2 flow rate set at 12 kg.Math.h.sup.1 and feed flow rate varying from 0.08 up to 0.55 kg.Math.h.sup.1.

(33) The extract is constituted of the totality of neutral lipids and the raffinate contains water and solid material (glucides, proteins, minerals, . . . ). The separation of dry matter and water can be performed by ultrafiltration.