METHOD FOR BIOLOGICAL IN-SITU METHANATION OF CO2 AND H2 IN A BIOREACTOR

Abstract

The invention relates to a method for the biological in-situ methanation of CO.sub.2 and H.sub.2 in a bioreactor. The method includes feeding an organic substrate into the bioreactor wherein at least part of the organic substrate is converted to a biogas comprising methane and carbon dioxide by means of microorganisms. The organic substrate includes crude fiber and at least 0.15 kg of crude fiber per m.sup.3 bioreactor volume per day is fed into to the bioreactor. The bioreactor is operated at between about 20-45° C. H.sub.2 is fed to the CO.sub.2 into the bioreactor to produce methane.

Claims

1. A method for the biological in-situ methanation of carbon dioxide and hydrogen in a bioreactor containing microorganisms and having a bioreactor volume, the method comprising: a) feeding an organic substrate to the microorganisms in the bioreactor to produce a biogas comprising methane and carbon dioxide, wherein the organic substrate includes crude fiber and organic dry substance (VS), and wherein the organic substrate fed to the bioreactor includes at least 0.15 kg of crude fiber per m.sup.3 of the bioreactor volume per day; b) operating the bioreactor at between 20-45° C.; c) feeding hydrogen into the bioreactor, wherein and at least part of the hydrogen together with the carbon dioxide is converted to methane by the microorganisms; and d) removing a product gas from the bioreactor, wherein the product gas includes methane.

2. The method according to claim 1 wherein the product gas further comprises carbon dioxide, hydrogen, and water vapor.

3. The method according to claim 1, wherein the bioreactor is operated at an organic loading rate of at least 2.5 kg VS per m.sup.3 of the bioreactor volume per day.

4. The method according to claim 1, wherein the bioreactor is operated at an organic loading rate of at least 3.0 kg VS per m.sup.3 of the bioreactor volume per day.

5. The method according to claim 1, wherein at least 0.2 kg crude fiber per m.sup.3 of the bioreactor volume per day is fed to the bioreactor,

6. The method according to claim 1, wherein at least 1.0 kg crude fiber per m.sup.3 of the bioreactor volume per day is fed to the bioreactor.

7. The method according to claim 1, wherein the hydrogen is fed into the bioreactor at a rate of at least 0.017 Nm.sup.3 per m.sup.3 of the bioreactor volume per hour.

8. The method according to claim 1, wherein the hydrogen is fed into the bioreactor at a rate of at least 0.125 Nm.sup.3 per m.sup.3 of the bioreactor volume per hour.

9. The method according to claim 1, further comprising feeding carbon dioxide to the bioreactor.

10. The method according to claim 9, further comprising adjusting the amount of carbon dioxide fed to the bioreactor such that the concentration of carbon dioxide in the product gas without the water vapor does not fall below a value of 5% (v/v).

11. The method according to claim 9, further comprising adjusting the amount of carbon dioxide fed to the bioreactor such that the concentration of carbon dioxide in the product gas without the water vapor does not fall below a value of 40% (v/v).

12. The method according to claim 9, further comprising regulating the concentration of carbon dioxide in the product gas by adjusting the amount of carbon dioxide fed to the bioreactor.

13. The method according to claim 9, further comprising regulating the concentration of carbon dioxide in the product gas by adjusting the amount of hydrogen fed to the bioreactor.

14. The method according to claim 9, further comprising feeding the carbon dioxide in an amount which adjusts the pH in the bioreactor to below 8.2.

15. The method according to claim 1, wherein the bioreactor has a NH.sub.4—N content and the method further comprises adjusting the NH.sub.4—N content of the bioreactor to levels below 6,000 mg/kg by feeding a low ammonium liquid and/or by the selection of the organic substrate.

16. The method according to claim 15, wherein the low-ammonium liquid is effluent from the bioreactor whose NH.sub.4—N content has been reduced via ammonium stripping.

17. The method according to claim 2, further comprising separating hydrogen from the product gas and returning the separated hydrogen to the bioreactor.

18. The method according to claim 2, further comprising separating carbon dioxide from the product gas and returning at least a portion of the separated carbon dioxide to the bioreactor.

19. A method for the biological in-situ methanation of carbon dioxide and hydrogen in a bioreactor containing microorganisms and having a bioreactor volume, the method comprising: a) feeding an organic substrate to the microorganisms in the bioreactor to produce a biogas comprising methane and carbon dioxide; b) operating the bioreactor at between 20-45° C.; c) feeding hydrogen to the carbon dioxide in the bioreactor to produce methane; d) removing a product gas from the bioreactor, wherein the product gas includes methane, carbon dioxide, and hydrogen; e) separating hydrogen from the product gas and returning the separated hydrogen to the bioreactor; and f) separating carbon dioxide from the product gas and returning at least a portion of the separated carbon dioxide to the bioreactor.

20. A method for the biological in-situ methanation of carbon dioxide and hydrogen in a bioreactor containing microorganisms and having a bioreactor volume, the method comprising: a) feeding an organic substrate to the microorganisms in the bioreactor to produce a biogas comprising methane and carbon dioxide, wherein the organic substrate includes crude fiber, and wherein the organic substrate fed to the bioreactor includes at least 0.15 kg of crude fiber per m.sup.3 of the bioreactor volume per day; b) operating the bioreactor at between 20-45° C.; c) feeding hydrogen to the carbon dioxide in the bioreactor to produce methane; d) feeding carbon dioxide into the bioreactor; e) removing a product gas from the bioreactor, wherein the product gas includes methane produced in step (a) and step (c), carbon dioxide, and hydrogen; f) adjusting the amount of carbon dioxide fed to the bioreactor such that the concentration of carbon dioxide in the product gas without the water vapor does not fall below a value of 10% (v/v).

Description

BRIEF DESCRIPTION OF THE FIGURES

[0053] The invention is explained in more detail below using four embodiments and associated drawings.

[0054] FIG. 1, FIG. 2 and FIG. 3 describe the schematic process sequence of some embodiments of the method according to aspects of the invention.

[0055] FIGS. 4 (a) and 4 (b) show diagrams of experimental data.

EMBODIMENT 1

[0056] FIG. 1 shows a schematic representation of the method in one embodiment with the feeding of CO.sub.2 from an ethanol plant into the bioreactor as well as an effluent treatment with ammonium stripping and recirculation of the resulting low-ammonium process liquid to adjust the NH.sub.4—N content in the bioreactor. Table 1 shows liquid and solid flows, such as the mass flows of organic substrate in the form of the organic loading rate, crude fiber, and low-ammonium process liquid. Table 2 shows the volume flow rates of the gases fed to and discharged from the bioreactor.

[0057] In this and the following examples, the volume flow rates of H.sub.2, CO.sub.2 and H.sub.2S leaving the bioreactor are referred to as H.sub.2, CO.sub.2, and H.sub.2S output rates, respectively, and include the volume of the respective gas under standard conditions (N for short, 273.15 K and 1.01325 bar) in m.sup.3 per m.sup.3 bioreactor volume discharged from the bioreactor per hour on average. This results in the unit Nm.sup.3/m.sup.3/h. The term “stillage” refers to the residue from the distillation of a grain mash containing ethanol. The term “whole stillage” is used as a synonym of stillage. In the context of the present invention, “wet cake” refers to the solid phase separated from the stillage by solid-liquid separation.

Step 1

[0058] A daily amount of 11,207 kg of whole stillage from a bioethanol plant is fed as organic substrate to a bioreactor with 1,000 m.sup.3 bioreactor volume. The TS content of the whole stillage corresponds to 0.19 kg TS per kg of whole stillage. The VS content of the whole stillage corresponds to 0.91 kg VS per kg TS of whole stillage. This results in a daily feed of whole stillage of 1,937 kg VS or, based on the bioreactor volume of the bioreactor, an organic loading rate of 1.94 kg VS/m.sup.3/d.

[0059] Based on Weender's food and feed analysis, the crude fiber content of the whole stillage corresponds to 0.039 kg crude fiber per kg TS. Calculated on the basis of the TS content of the whole stillage and the mass flow of whole stillage fed, the result is a very low supply of crude fiber via the whole stillage of only 0.08 kg crude fiber/m.sup.3/d.

[0060] Surprisingly, it was found that a significantly higher crude fiber content favors the CH.sub.4 production rate from fed H.sub.2, so 4,065 kg/d of wet cake is fed to the bioreactor in addition to the whole stillage. Wet cake contains a high TS content of approx. 0.33 kg TS per kg of wet cake, and an VS content of 0.96 kg VS per kg TS and is typically sold as feed. With a high crude fiber content of 0.14 kg crude fiber per kg TS, wet cake is not a typical organic substrate for biogas production. However, in this embodiment example, it proves to be very useful to significantly increase the crude fiber supply by 0.19 kg crude fiber/m.sup.3/d.

[0061] Moreover, 6000 kg/d of purified effluent from step 3, 15.8 Nm.sup.3 CO.sub.2/h from a bioethanol plant and 25 Nm.sup.3 H.sub.2/h from an electrolysis plant are fed to the bioreactor.

TABLE-US-00001 TABLE 1 Liquid and solid flows Embodiment 1 Numeric value Unit Organic loading rate 1.9 kg VS/m.sup.3/d Whole stillage Organic loading rate Wet cake 1.3 kg VS/m.sup.3/d Crude fiber Whole stillage 0.08 kg crude fiber/m.sup.3/d Crude fiber Wet cake 0.19 kg crude fiber/m.sup.3/d Low-ammonium process 6 kg/m.sup.3/d liquid from step 3

Step 2

[0062] Organic substrate is converted in the bioreactor to the gases CH.sub.4, CO.sub.2 and H.sub.2S, among others. The H.sub.2 supplied from the electrolysis is converted to approx. 60%. The remaining H.sub.2 leaves the bioreactor with the other gases via the product gas (see Table 2).

[0063] The supply of CO.sub.2 from the bioethanol plant results in the CO.sub.2 content in the product gas being approx. 45% (v/v) based on product gas without the water content. Without the CO.sub.2 supply from the bioethanol plant, the CO.sub.2 content in the product gas would only be approx. 37% (v/v) based on product gas without the water content, which would lead to an unwanted, higher pH value. The supply of CO.sub.2 from the bioethanol plant thus results in a more stable reaction in the bioreactor.

[0064] The temperature in the bioreactor is controlled at 37° C.

[0065] In the embodiment example described here, a CH.sub.4 production rate of 0.053 Nm.sup.3/m.sup.3/h would be achieved if no H.sub.2 and CO.sub.2 were fed. Using the invention described here, the CH.sub.4 production rate increases to 0.056 Nm.sup.3/m.sup.3/d by feeding and converting H.sub.2 and CO.sub.2. The CH.sub.4 production rate from fed H.sub.2 corresponds to 0.0038 Nm.sup.3/m.sup.3/h. This increases the utilization of the available bioreactor volume.

TABLE-US-00002 TABLE 2 Gas flows Embodiment 1 Numeric value Unit Gas feed rates H.sub.2 from electrolysis 0.025 Nm.sup.3/m.sup.3/h CO.sub.2 from bioethanol plant 0.016 Nm.sup.3/m.sup.3/h Product gas CH.sub.4 production rate 0.056 Nm.sup.3/m.sup.3/h CO.sub.2 output rate 0.055 Nm.sup.3/m.sup.3/h H.sub.2 output rate 0.010 Nm.sup.3/m.sup.3/h H.sub.2S output rate 0.001 Nm.sup.3/m.sup.3/h

Step 3

[0066] The effluent of the bioreactor is fed into an effluent treatment using ammonium stripping, and ammonium is largely removed. The resulting process liquid, which is low in ammonium, has an NH.sub.4—N content of 500 mg/kg. A part of this low-ammonium process liquid is fed to the bioreactor to adjust the NH.sub.4—N content, thereby adjusting the NH.sub.4—N content to below 6000 mg/kg.

EMBODIMENT 2

[0067] Another option for process control consists of utilizing the CO.sub.2 produced from organic substrate during biogas production instead of CO.sub.2 from a bioethanol plant. For this purpose, a product gas treatment is added to the method described in embodiment 1, in which CO.sub.2 is separated from the product gas at an hourly circulation rate of 0.016 Nm.sup.3 CO.sub.2 per m.sup.3 bioreactor volume and fed back into the bioreactor. This CO.sub.2 replaces the CO.sub.2 flow from the bioethanol plant described in embodiment 1. FIG. 2 shows a schematic illustration of this method. Table 3 shows the gas flows. The liquid flows in this example are the same as in embodiment 1 and shown in Table 1. The temperature in the bioreactor is still controlled at 37° C. and the effluent treatment from step 3 in embodiment 1 is maintained.

[0068] In this context and as used hereinafter, circulation rate refers to the volume of the respective gas under standard conditions (N for short, 273.15 K and 1.01325 bar) in m.sup.3 per m.sup.3 bioreactor volume, which is partially separated from the product gas in the product gas treatment and fed back into the bioreactor on average per hour. This results in the unit Nm.sup.3/m.sup.3/h.

[0069] As in embodiment 1, the CO.sub.2 content in the product gas is 45% (v/v) based on product gas without the water component and thus advantageously higher than it would be at 37% (v/v) based on product gas without the water component and if CO.sub.2 had not been fed.

[0070] Since the changed CO.sub.2 source has no effect on the CH.sub.4 production rate with otherwise constant operating conditions as in embodiment 1, the CH.sub.4 production rate based on the supplied H.sub.2 is also 0.0038 Nm.sup.3/m.sup.3/h.

TABLE-US-00003 TABLE 3 Gas flows Embodiment 2 Numeric value Unit Gas feed rates H.sub.2 from electrolysis 0.025 Nm.sup.3/m.sup.3/h CO.sub.2 from product gas treatment 0.016 Nm.sup.3/m.sup.3/h Product gas CH.sub.4 0.056 Nm.sup.3/m.sup.3/h CO.sub.2 output rate 0.055 Nm.sup.3/m.sup.3/h H.sub.2 output rate 0.010 Nm.sup.3/m.sup.3/h H.sub.2S output rate 0.001 Nm.sup.3/m.sup.3/h Product gas treatment CO.sub.2 circulation rate 0.016 Nm.sup.3/m.sup.3/h Processed product gas CH.sub.4 production rate 0.056 Nm.sup.3/m.sup.3/h CO.sub.2 0.039 Nm.sup.3/m.sup.3/h H.sub.2 0.010 Nm.sup.3/m.sup.3/h H.sub.2S 0.001 Nm.sup.3/m.sup.3/h

EMBODIMENT 3

[0071] If feeding product gas into the natural gas network is planned, certain rules stipulated by the network operator regarding the composition of the gas must be observed. As a rule, for example, the CH.sub.4 content must be greater than 95% (v/v), the H.sub.2 content must be less than 2% (v/v) and the H.sub.2S content must be approx. 0% (v/v) based on gas without the water content.

[0072] FIG. 3 shows a schematic representation of a modified method according to embodiment 2 with upgraded product gas treatment and gas recirculation into the bioreactor. H.sub.2S is almost completely removed from the product gas. A large part of the CO.sub.2 is removed from the product gas and some of it is returned to the bioreactor, while the rest is sent for other utilization. H.sub.2 is separated via a membrane. In terms of investment and operating costs, the fact that only one separation stage is used proves to be advantageous. This is sufficient to maintain a typical H.sub.2 maximum concentration of 2% (v/v) for gas injection into a natural gas grid, but results in some CH.sub.4 being returned to the bioreactor along with the H.sub.2.

[0073] Table 4 shows the gas flows in this embodiment. The liquid flows in this example are the same as in embodiment 1 and shown in Table 1. The temperature in the bioreactor is still controlled at 37° C. and the effluent treatment from step 3 in embodiment 1 is maintained.

[0074] Recirculation of H.sub.2 increases the retention time of H.sub.2 in the bioreactor. This allows the conversion of the H.sub.2 coming from the electrolyzer to be increased from 60% to 95% (v/v). The CH.sub.4 production rate from supplied H.sub.2 thus corresponds to 0.00594 Nm.sup.3/m.sup.3/h. Due to the increased amounts of recycled gas, the increase in recycled CO.sub.2 is also necessary to keep the CO.sub.2 content in the product gas high. Thus, the advantages of a stable CO.sub.2 concentration in the product gas described in embodiment 1 can be maintained.

TABLE-US-00004 TABLE 4 Gas flows Embodiment 3 Numeric value Unit Gas feed rates H.sub.2 feed rate (from electrolysis) 0.025 Nm.sup.3/m.sup.3/h CO.sub.2 from product gas treatment 0.028 Nm.sup.3/m.sup.3/h H.sub.2 feed rate (from product gas 0.015 Nm.sup.3/m.sup.3/h treatment) CH.sub.4 from product gas treatment 0.004 Nm.sup.3/m.sup.3/h Product gas CH.sub.4 0.063 Nm.sup.3/m.sup.3/h CO.sub.2 output rate 0.065 Nm.sup.3/m.sup.3/h H.sub.2 output rate 0.016 Nm.sup.3/m.sup.3/h H.sub.2S output rate 0.001 Nm.sup.3/m.sup.3/h Product gas treatment CO.sub.2 circulation rate 0.028 Nm.sup.3/m.sup.3/h CO.sub.2 for other utilization 0.034 Nm.sup.3/m.sup.3/h H.sub.2 circulation rate 0.015 Nm.sup.3/m.sup.3/h CH.sub.4 circulation rate 0.004 Nm.sup.3/m.sup.3/h H.sub.2S 0.001 Nm.sup.3/m.sup.3/h Processed product gas CH.sub.4 0.058 Nm.sup.3/m.sup.3/h production rate CO.sub.2 0.002 Nm.sup.3/m.sup.3/h H.sub.2 0.001 Nm.sup.3/m.sup.3/h H.sub.2S 0.000 Nm.sup.3/m.sup.3/h

EMBODIMENT 4

[0075] Embodiment 4 includes experiments on a pilot-plant scale.

Method

[0076] Two bioreactors “A” and “B” of the stirring-tank-reactor type were fed with different organic substrates as well as H.sub.2 and CO.sub.2.

[0077] Bioreactor A was fed with whole stillage and low-ammonium process liquid from a bioethanol plant. Feeding was carried out semi-continuously, distributed in five equal intervals per day.

[0078] Bioreactor B was fed with cereal straw and low-ammonium process liquid. Feeding was done once a day.

[0079] The bioreactor volume corresponded to 75 l (bioreactor A) and 55 l (bioreactor B). Mixing was performed with central stirrers at a rotational speed of 370 rpm for both bioreactors. The temperature of the bioreactor volume was controlled using water as a heating medium at about 39° C., which flowed around the bioreactors via a double casing. The supplied quantities of H.sub.2 and CO.sub.2 were controlled separately for each gas via thermal mass flow controllers (EL-Flow Select, Bronkhorst Deutschland Nord GmbH). Product gas volume flows were determined in both bioreactors using drum gas meters (TG 0.5, Dr.-Ing. RITTER Apparatebau GmbH & Co. KG). The gas composition of the product gas was determined once a week using gas chromatography (MobilGC, ECH Elektrochemie Halle GmbH; configuration: Hayesep QS column 45° C., molecular sieve column 55° C., thermal conductivity detector, carrier gas argon). The TS content of the substrates was determined gravimetrically after drying at 105° C. until mass constancy was achieved. The ash content was determined gravimetrically after annealing at 650° C. for at least two hours. The VS content was calculated based on the difference between the dry substance content and the ash content. The crude fiber content was determined according to the VDLUFA method book (see section Definitions). The low-ammonium process liquid was neglected in the determination of the organic loading rate and the fed crude fiber.

Results and Discussion

[0080] FIG. 4 shows data from the two bioreactors dated Apr. 16, 2021 to May 19, 2021. Prior to this, both reactors had already been operated for more than one year, hence transient effects of the start-up phase can be excluded and neglected in the interpretation of the data.

[0081] In both bioreactors, organic substrate is converted to biogas by means of microorganisms, as indicated by a lower amount of VS in the effluent compared to the reactor feed, as well as by means of the measured CH.sub.4 production rates, which are higher than the calculated CH.sub.4 production rates from fed H.sub.2 and therefore must originate from organic substrate decomposition (data not shown).

[0082] The organic loading rate is comparable for both reactors at approx. 3 kg VS/m.sup.3/d. However, the feeding of crude fiber differs significantly. While in bioreactor A only about 0.14 kg crude fiber/m.sup.3/d was fed via the whole stillage, the amount in bioreactor B via the straw was significantly higher at about 1.2 kg crude fiber/m.sup.3/d.

[0083] Bioreactor B exhibited significantly higher performance for in-situ methanation of H.sub.2 and CO.sub.2. Even at high H.sub.2 feed rates of 0.197 Nm.sup.3/m.sup.3/h, low H.sub.2 output rates of about 0.014 Nm.sup.3/m.sup.3/h were measured. Thus, according to equation (2), the CH.sub.4 production rate from fed H.sub.2 is about 0.046 Nm.sup.3/m.sup.3/h. In bioreactor A, on the other hand, higher H.sub.2 output rates were measured than in bioreactor B of 0.048 Nm.sup.3/m.sup.3/h, despite significantly lower H.sub.2 feed rates of only 0.081 Nm.sup.3/m.sup.3/h. This corresponds to a CH.sub.4 production rate from fed H.sub.2 in bioreactor A of 0.008 Nm.sup.3/m.sup.3/h.

LIST OF REFERENCES

[0084] Schiraldi C., De Rosa M. (2014) Mesophilic Organisms. In: Drioli E., Giorno L. (eds) Encyclopedia of Membranes. Springer, Berlin, Heidelberg. https://doi.org/10.1007/978-3-642-40872-4_1610-2