METHOD FOR PRODUCING HIGHLY PURE HYDROGEN BY COUPLING PYROLYSIS OF HYDROCARBONS WITH ELECTROCHEMICAL HYDROGEN SEPARATION

Abstract

The present invention comprises a process for producing hydrogen, wherein in a first stage hydrocarbons are decomposed into solid carbon and into a hydrogen-containing gaseous product mixture, the hydrogen-containing gaseous product mixture, which has a composition in respect of the main components CH4, N2, and H2 of 20% to 95% by volume H2 and 80% to 5% by volume CH4 and/or N2, is discharged from the first stage at a temperature of 50 to 300° C., and this is supplied at a temperature differing from this exit temperature by not more than 100° C. to an electrochemical separation process and, in this second stage, the hydrogen-containing product mixture is separated in the electrochemical separation process at a temperature of 50 to 200° C. into hydrogen having a purity of >99.99% and a remaining residual gas mixture.

Claims

1.-10. (canceled)

11. A process for producing hydrogen, wherein in a first stage hydrocarbons in a fixed-bed reactor, fluidized-bed reactor or moving-bed reactor in the presence of solid carrier materials having a granule size of 0.05 to 100 mm are decomposed into solid carbon and into a hydrogen-containing product mixture, the hydrogen-containing gaseous product mixture, which has a composition in respect of the main components CH.sub.4, N.sub.2, and H.sub.2 of 20% to 95% by volume H2 and 80% to 5% by volume CH.sub.4 and/or N.sub.2, is discharged from the first stage at a temperature of 50 to 300° C., wherein the cooling of the hot product streams is used to heat the feed streams, and this is supplied at a temperature differing from this exit temperature by not more than 100° C. to an electrochemical separation process and, in this second stage, the hydrogen-containing product mixture is separated in the electrochemical separation process at a temperature of 50 to 200° C. into hydrogen having a purity of >99.99% and a remaining residual gas mixture.

12. The process according to claim 11, wherein the electrochemical separation process in the second stage uses a membrane electrode assembly and the membrane is a polymer membrane selected from the group of sulfonated polyether ether ketones, sulfonated polybenzimidazoles, sulfonated fluorinated hydrocarbon polymers, perfluorinated polysulfonic acids, styrene-based polymers, poly(arylene ethers), polyimides, and polyphosphazenes.

13. The process according to claim 12, wherein polybenzimidazoles based on polybenzimidazole and phosphoric acid are used as polymer membranes.

14. The process according to claim 11, wherein the decomposition in the first stage is carried out at a temperature of 900° C. to 1200° C. for a residence time of 1 s to 1 min.

15. The process according to claim 11, wherein the cooling of the hot product-containing gas from reaction temperature to an exit temperature of 50° C. to 300° C. takes place in a solid bed.

16. The process according to claim 11, wherein 90 to 99.99% of the amount of residual gas remaining from the electrochemical separation process is recirculated to the first stage.

17. The process according to claim 11, wherein the composition in respect of the main components CH.sub.4, N.sub.2, and H.sub.2 is from 80% to 90% by volume H2 and 20% to 10% by volume CH.sub.4 and/or N.sub.2.

18. The process according to claim 11, wherein no catalyst is present in the first stage.

19. The process according to claim 11, wherein the pyrolysis product gas comprises more than 3% CO.

20. The process according to claim 11, wherein both stages are carried out at an absolute pressure of 1 bar to 30 bar and the pressure difference between the two stages is within a range from 0.001 bar to 5 bar.

21. The process according to claim 11, wherein the energy required for the decomposition reaction in the first stage is provided autothermally or via low-temperature plasma.

22. The process according to claim 11, wherein the autothermal pyrolysis process comprises the following steps: 1) Providing a particle bed composed of a carrier material. 2) Burning a reactant gas or product gas with air to produce a hot pyrolysis gas for providing the enthalpy of reaction. 3) Mixing this hot pyrolysis gas with the reactant gas, such that the reactant gas pyrolyzes to H2 and carbon. 4) Contacting the particle bed of carrier material with the carbon-containing hot pyrolysis gas, such that the particle bed of carrier material is heated and carbon is deposited in the particle bed. 5) Passing cold reactant gas over this heated particle bed, such that the reactant gas is heated and the carbon-laden particle bed is cooled. 6) Replacing the cooled, laden particle bed with a cold particle bed.

23. The process according to claim 11, wherein the low-temperature plasma pyrolysis process comprises the following steps: 1) Providing a particle bed composed of a carrier material. 2) Contacting the particle bed of carrier material with the carbon-containing hot pyrolysis gas, such that the particle bed of carrier material is heated and carbon is deposited in the particle bed. 3) Passing cold feed gas consisting of reactant gas and recirculated gas over this heated particle bed, such that this feed gas is heated and the carbon-laden particle bed is cooled. 4) Further heating of the feed gas by a plasma burner to produce the hot pyrolysis gas. 5) Replacing the cooled, laden particle bed with a cold particle bed.

24. The process according to claim 11, wherein the hydrogen present after the electrochemical separation process is supplied to a hydrogen car.

Description

EXAMPLES

[0154] The process examples were calculated with the aid of the company's thermodynamic simulator Chemasim, which is analogous to Aspen.sup.+. The reactor design was executed in Excel on the basis of thermodynamic simulation.

[0155] The process was by way of example calculated for a H2 capacity of 1000 kg/day, or 42 kg/h.

[0156] The value is based on the largest H2 filling stations currently under discussion.

[0157] As a measure for comparison purposes, the following process parameters are employed: [0158] As a measure for variable costs: [0159] the natural gas/methane requirement [0160] the electricity consumption [0161] the generation of pyrolytic carbon as a credit [0162] As a measure for capital costs: [0163] the heat-transfer capacity, or transferred specific heat based on amount of product, that is relevant to capital costs. [0164] In addition, the specific investment costs as stated in today's literature are used. [0165] As a measure of ecology, which is of course the main driver for the development of H2 filling stations: [0166] the carbon footprint of the process including the carbon footprint of the required grid electricity.

[0167] For operation of an H2 filling station, the sole practical option is grid electricity, because operation needs to be available around the clock, irrespective of the weather and the position of the sun.

[0168] For grid electricity, the future electricity mix forecast for 2030 for the EU 27 was used, which comprises 19% nuclear, 33% fossil, and 48% renewable energy. The data are taken from [7] and represent a European average. This results in a calculated carbon footprint of 190 kg CO2/MWh.sub.el. for the electricity mix in the EU 27 in 2030.

[0169] It is also assumed that the filling station is connected to a 25 bar natural gas network.

[0170] For the process comparison, all processes produce 20 bar of H2.

[0171] All examples are calculated assuming zero losses.

Prior Art

[0172] Electrolysis:

[0173] The comparison of the inventive processes with the prior art uses electrolysis performance data as published for alkaline electrolysis in [8].

[0174] According to these data, the efficiency of the overall system operating at atmospheric pressure and 80° C. is 68%. This corresponds to a specific electrical energy consumption of 48.4 kWh/kg H2. If the H2 is compressed from 1 bar to 20 bar, another 1.6 kWh/kg H2 is required.

[0175] The specific electrical energy requirement is therefore 50.0 kWh.sub.el/kg H2 in total.

[0176] The specific carbon footprint is then 9.50 kg CO2/kg H2.

[0177] 32% (=100%−68%) of the electrical energy is converted into heat and must be dissipated into the environment via heat-exchanger surfaces. The specific heat-transfer energy is herewith 22.8 kWh/kg H2.

[0178] According to [4], the specific investment cost is €3070 a/t H2.

[0179] The intermediate cooling for the H2 compression from 1 to 20 bar requires 1.7 kWh/kg H2.

[0180] Specifically, a total of 22.8+1.7=24.5 kWh of heat is thus transferred per kg of H2.

[0181] The electrolysis requires herewith per kg of H2: [0182] 50.0 kWh of current [0183] 24.5 kWh of heat transferrer

[0184] and produces per kg of H2: [0185] 9.50 kg of CO2

[0186] Mini-SMR:

[0187] Detailed data for a mini-SMR unit for 90 kg/h H2 are reported in [9]. This corresponds to only twice the capacity of a 42 kg H2/h capacity serving here as a basis for an H2 filling station and is therefore very well suited for the comparison of the prior art with the inventive variants. An update of the cost data based on the process data and list of machines and equipment in [9] is given in [10].

[0188] According to this, 3.1 kg of natural gas and 2.1 kWh.sub.el of electricity are required per kg of H2. The natural gas here has the following composition in % by weight: 88.7% CH4, 4.7% C2H6, 3.9% C3H8, 1.3% N2, and 1.3% CO2. The reported electricity requirement covers not just the actual requirement of the process, but also the compression of the natural gas from 7 to 22 bar prior to the process and compression of the H2 from 21 bar to 207 bar after the process. For the actual process, the reported data give rise to an electricity requirement of 0.2 kWh.sub.el/kg H2.

[0189] 0.04 kg CO2/kg H2 results from the grid electricity supplied and 8.42 kg CO2/kg H2 results from production. In total, the production of 1 kg of H2 according to mini-SMR technology prior art thus produces 8.46 kg of CO2.

[0190] The reported values for the heat transferrers give rise to a calculated value of 18.5 kW per kg H2/h for installed specific heat-transfer capacity. However, this does not include the cost-relevant heat-transfer capacity of the reformer. No information on this was provided. According to [9], the specific investment cost of a mini-SMR is € 12 100 a/t H2.

[0191] To achieve better comparability of the prior art with the inventive variants, the SMR process published in [9] was recalculated using the company's thermodynamic process simulator Chemasim, which was also used for the calculation of the inventive variants. The unit ratios, operating parameters, and heat-transfer capacities reported in [9] were specified and the unknown heat-transfer capacity for the reformer thus determined. This gives a value of 8.9 kW per kg H2/h. The total specific heat-transfer capacity is thus 18.5+8.9=27.4 kW per kg H2/h.

[0192] If 100% CH4 is used in the calculation instead of a natural gas having the composition stated above, the unit ratios and specific heat-transfer capacities change only marginally. For the sake of simplicity, the following results are therefore based on simulations with 100% CH4 as the feed gas.

[0193] The mini-SMR requires herewith per kg of H2: [0194] 3.10 kg of CH4 [0195] 0.2 kWh of current [0196] 27.4 kWh of heat transferrer

[0197] and produces per kg of H2: [0198] 8.46 kg of CO2

[0199] Tube Trailer H2:

[0200] Detailed data for a world-scale SMR plant for 9058 kg/h H2 are reported in [11]. According to the “Major Equipment” list, 847.4 MMBTU/h of heat is—as in the case of the mini-SMR—transferred therefor, which equates to 27.4 kWh/kg H2. However, more natural gas is required in the case of the world-scale plant than in the case of the mini-SMR. The world-scale plant accordingly requires 3.32 kg of natural gas/kg H2 instead of 3.10 kg of natural gas/kg H2 in the case of the mini-SMR. On the other hand, the world-scale plant does however also produce, in addition to H2, 4.4 kg of steam/kg H2. Since the appraisal of steam depends very much on local conditions, it was assumed here for the sake of simplicity that the world-scale plant has the same unit ratios as the mini-SMR. The cost advantage of a world-scale plant over a mini-SMR lies in economy of scale.

[0201] For transport, it is assumed that the H2 is transported by road to the filling stations in 500 bar containers on trailers and that these containers are emptied to 21 bar at the filling station before being transported back to the world-scale plant. For transport, the H2 must be compressed from 20 to 500 bar at the world-scale plant [5]. This requires the use of 1.6 kWh/kg H2. This high pressure in the containers is however advantageous when compressing at the filling station to the final pressure of e.g. 950 bar for refueling cars. However, given that the pressure decreases from 500 to 20 bar at the filling station during emptying of the containers, an average initial pressure in the container of (500+20)/2=260 bar is assumed for the comparison with the other variants. This reduces by 1.3 kWh/kg the energy needed for further compression from 260 bar to e.g. 950 bar compared to compression from 20 bar to 950 bar. H2 transport thus ultimately requires the use of 1.6−1.3=0.3 kWh/kg H2 more electricity than in the case of the mini-SMR.

[0202] According to [5], the specific investment cost for compression is € 430 a/t H2.

[0203] With 500 bar containers, a maximum of 1344 kg of H2 can currently be transported by a 40 t tank truck [5] that is emptied at the filling station down to a pressure of 21 bar. This then leaves behind 54 kg of H2 in the containers, which is returned to the world-scale plant. According to [5], the specific investment cost for the additional expenditure on storage at the filling station is € 4740 a/t H2. The total specific investment cost is then € 430+4740=5170 a/t H2.

[0204] It is further assumed that a tank truck of this kind requires 35 liters of diesel per 100 km. In terms of energy content, this corresponds to 31 kg of natural gas per 100 km and currently represents a best figure.

[0205] For 42 kg/h of H2, 0.78 journeys per day are necessary (=42*24/(1344−54)). If the energy consumption and associated CO2 emissions are calculated for different distances—e.g. for 100 km and for 500 km—between the centralized world-scale plant and the decentralized filling station, the following results are obtained: The tube trailer variant requires per kg of H2: [0206] 3.10 kg of CH4 for generation in the centralized world-scale unit [0207] +0.08 kg of CH4 equivalents in the form of diesel for a distance of 100 km [0208] +0.30 kg of CH4 equivalents in the form of diesel for a distance of 500 km [0209] 0.6 kWh of current [0210] 27.4 kWh of heat transferrer

[0211] and produces per kg of H2: [0212] 8.75 kg of CO2 for a distance of 100 km and [0213] 9.18 kg of CO2 for a distance of 500 km and

[0214] Inventive Process Variants

[0215] ATP&EHS: (FIG. 4)

[0216] An example calculation was carried out for a combination of an autothermally operated pyrolysis (ATP) and an EHS. The advantage of this combination lies in a low consumption of electricity and natural gas and also a low heat-transfer capacity. 42 kg/h of high-purity H2 is generated.

[0217] Natural gas is prepurified. This can take place for example by catalytic desulfurization, as described in [9].

[0218] 147 kg/h of purified natural gas at a pressure level of 25 bar is supplied to the process at ambient temperature (25° C.).

[0219] 25 kg/h thereof is withdrawn to be burnt in a burner with air to hot burner gas to cover the energy requirement. 430 kg/h of burner air is compressed from ambient pressure and temperature to 1.5 bar. This needs 5 kW of electric power.

[0220] The rest of the natural gas is mixed with 227 kg/h of recirculated gas in a jet pump. In the jet pump, the initial pressure of the natural gas is used to compress the recirculated gas from 1.0 to 1.5 bar.

[0221] The feed gas enters the particle bed of segment (2c) at a temperature of 28° C. and is heated to 1000° C. therein. This is accompanied by cooling of the particle bed. A temperature front develops, which moves from bottom to top. This is accompanied by a thermal transfer of 199 kW.

[0222] During this heating, part of the natural gas already undergoes pyrolysis.

[0223] After exiting segment (2c), the gas is mixed with the hot burner gas and further reactions commence.

[0224] Although no catalyst is present, it must be assumed that part of the natural gas will be reformed to CO and H2 with the water that is formed during combustion. It is assumed here by way of example that 10% of the combustion water reacts. It is also assumed that the CO2 from the combustion reacts in the particle bed in segment (2b) with the pyrolytic carbon formed to form CO according to the Boudouard equilibrium.

[0225] In segment (2b), the hot reaction gas heats the particle bed and is at the same time itself cooled. This is similarly accompanied by a thermal transfer of 199 kW.

[0226] The product gas (757 kg/h) cooled to 160° C. comprises 15 mol % of CO. This CO is in a WGS reaction with steam converted to CO2 and H2 down to a residual concentration of 0.2%. The reaction gives rise to 69 kW of excess heat, which must be dissipated. This is done by generating 5 bar of steam, which is needed as additional steam (93 kg/h) for the WGS reaction.

[0227] In the EHS, 99% of the H2 formed is separated electrochemically from the product gas of the WGS (850 kg/h). This needs 177 kW of electric power.

[0228] The residual anode offgas (808 kg/h) is split into the recirculated gas that is recycled into the process and the offgas that is burned in the flare, thereby generating 232 kg CO2/h. 42 kg/h H2 exits the EHS at a pressure of 1 bar. The compression to 20 bar needs 68 kW.sub.el of electric power. 65 kW must be abstracted from the intermediate cooling as heat flows. In order to be able to provide regenerative heat-transfer capacity of 199 kW in segments (2b) and (2c), 1034 kg/h of fresh pyrolytic carbon must be introduced into segment (2a). 1080 kg/h of pyrolytic carbon is withdrawn from segment (2d). The difference, 46 kg/h, is generated as pyrolytic carbon product.

[0229] The ATP&EHS process produces herewith per kg of H2: [0230] 1.1 kg of high-purity pyrolytic carbon

[0231] and requires therefor per kg of H2: [0232] 3.5 kg of CH4 [0233] 6.0 kWh of current [0234] 10.0 kWh of heat transferrer

[0235] and produces per kg of H2: [0236] 6.7 kg of CO2

[0237] LT Plasma&EHS: (FIG. 5)

[0238] An example calculation was carried out for a combination of a pyrolysis operated with a low-temperature plasma (LT plasma) and an EHS. The plasma here has the role of increasing the rate of reaction. The advantage of this combination lies in the relatively low process temperatures and the absence of CO, which has a beneficial effect on the energy requirement of the EHS.

[0239] 42 kg/h of high-purity H2 is generated.

[0240] Natural gas is prepurified. This can take place for example by catalytic desulfurization, as described in [9].

[0241] 168 kg/h of purified natural gas at a pressure level of 25 bar is supplied to the process at ambient temperature (25° C.) and mixed with 205 kg/h of recirculated gas in a jet pump. In the jet pump, the initial pressure of the natural gas is used to compress the recirculated gas from 1.0 to 1.5 bar.

[0242] The feed gas (373 kg/h) enters the particle bed of segment (2c) at a temperature of 28° C. and is heated to 700° C. therein. This is accompanied by cooling of the particle bed. A temperature front develops as a result, which moves from bottom to top. This is accompanied by a thermal transfer of 244 kW.

[0243] After exiting segment (2c), the gas molecules are excited in a low-temperature plasma device, for example by means of pulsed microwaves, and then passed into segment (2b). In segment (2b), the hot reaction gas heats the particle bed and is at the same time itself cooled to 160° C. This is similarly accompanied by a thermal transfer of 244 kW.

[0244] The product gas (248 kg/h) cooled to 160° C. is passed into an EHS in which 91% of the H2 formed is separated electrochemically from the product gas. This needs 102 kW of electric power.

[0245] 1 kg/h from the residual anode offgas (206 kg/h) is withdrawn as a purge stream in order to prevent accumulation of inert components.

[0246] 42 kg/h H2 exits the EHS at a pressure of 1 bar. The compression to 20 bar needs 68 kW.sub.el of electric power. 65 kW must be abstracted from the intermediate cooling as heat flows. In order to be able to provide regenerative heat-transfer capacity of 244 kW in segments (2b) and (2c), 1828 kg/h of fresh pyrolytic carbon must be introduced into segment (2a). 1953 kg/h of pyrolytic carbon is withdrawn from segment (2d). The difference, 125 kg/h, is generated as pyrolytic carbon product.

[0247] The LT plasma&EHS process produces herewith per kg of H2: [0248] 3.0 kg of high-purity pyrolytic carbon

[0249] and requires therefor per kg of H2: [0250] 4.0 kg of CH4 [0251] 10.4 kWh of current [0252] 9.4 kWh of heat transferrer

[0253] and produces per kg of H2: [0254] 2.0 kg of CO2

[0255] RH Pyrolysis&EHS: (FIG. 6)

[0256] An example calculation was carried out for a combination of an electrically heated pyrolysis in which the pyrolytic carbon bed functions as resistance heating (RH pyrolysis) and an EHS. The principle of RH pyrolysis is described for example in U.S. Pat. No. 2,982,622. The advantage of this combination lies in the simplicity of the reactor and in the higher possible operating pressures associated with its construction, which results in smaller reactor dimensions and a subsequently lower expenditure on compression for H2. The absence of CO moreover has a beneficial effect on the energy requirement of the EHS. The combination of the RH pyrolysis with the EHS allows the construction of a gas circuit having a high proportion of H2. The higher H2 level reduces soot formation and additionally lowers the energy requirement in the EHS. In addition, the recirculation of gas opens up new possibilities for conveying the pyrolytic carbon circulation (for example pneumatic conveyance) and enables better heat integration

[0257] 42 kg/h of high-purity H2 is generated.

[0258] Natural gas is prepurified. This can take place for example by catalytic desulfurization, as described in [9].

[0259] 167 kg/h of purified natural gas at a pressure level of 25 bar is supplied to the process at ambient temperature (25° C.) and mixed with 68 kg/h of recirculated gas in a jet pump. In the jet pump, the initial pressure of the natural gas is used to compress the recirculated gas from 5.0 to 5.2 bar.

[0260] The feed gas (235 kg/h) enters the particle bed at the bottom of the fluidized-bed reactor at a temperature of 28° C. and is heated therein to 1000° C. In return, the pyrolytic carbon bed is cooled as it slips downwards. In this countercurrent heat exchange, there is a thermal transfer of 367 kW.

[0261] 888 kg/h of particles with a temperature of 28° C. is withdrawn from the bottom of the reactor via feeders. 763 kg/h is recirculated for heat integration (367 kW+315 kW) and fed back into the reactor at the top via feeders. 125 kg/h of pyrolytic carbon is withdrawn as a high-purity product.

[0262] In the reactor, an electric current is in the reaction zone conducted through the particle bed to cover the heat of reaction (262 kW.sub.el). The heat of reaction of 262 kW thus introduced does not influence the amount of pyrolytic carbon that needs to be recirculated and is therefore not material to capital costs. The carbon formed during methane cracking results in a growth of pyrolytic carbon particles.

[0263] The product gas from methane cracking flows upwards and heats the recirculated particles as they slip downwards. In return, the product gas is cooled. The degree of heat integration can be controlled by the amount of recirculating gas. In this countercurrent heat exchange, there is a thermal transfer of 315 kW.

[0264] The product gas (110 kg/h) cooled to 160° C. is passed into an EHS in which 50% of the H2 formed is separated electrochemically from the product gas. This needs 63 kW.sub.el of electric power.

[0265] The residual anode offgas (68 kg/h) is recirculated. To prevent accumulation of inert components, 0.1 kg/h from the recirculated gas is withdrawn.

[0266] 42 kg/h H2 exits the EHS at a pressure of 5 bar. The compression to 20 bar needs 29 kW.sub.el of electric power. 38 kW must be abstracted from the intermediate cooling as heat flows.

[0267] The RH pyrolysis&EHS process produces herewith per kg of H2: [0268] 3.0 kg of high-purity pyrolytic carbon

[0269] and requires therefor per kg of H2: [0270] 4.0 kg of CH4 [0271] 8.4 kWh of current [0272] 19.2 kWh of heat transferrer

[0273] and produces per kg of H2: [0274] 1.6 kg of CO2

[0275] Comparison:

[0276] Table 1 summarizes the results of the example calculations. The results show clearly that the inventive process concepts are able to produce H2 with a smaller carbon footprint than is possible according to the current state of the art. The smaller carbon footprint is the main driver for H2 mobility.

[0277] In addition, the heat-transfer capacities for the inventive process concepts that are relevant to capital costs are smaller than in the case of the current state of the art.

[0278] The increased consumption of natural gas in the case of the inventive process concepts is offset by the additional recovery of high-purity carbon. This increases the raw material yield and thus the added value.

[0279] According to the prior art, an SMR is in terms of mass able to process only about 32% (=1 kg H2/3.1 kg CH4) of the methane used in a value-adding manner. The inventive process concepts are however able to utilize up to 100% (=(1 kg H2+3 kg C)/4 kg CH4) of the methane used.

[0280] Moreover, the inventive process concepts have only one fifth to one ninth the power requirement of a water electrolysis. A low power requirement is however important, particularly with regard to the expansion in renewable energies that will be necessary in the future, since mobility is here in competition with other energy consumers.

TABLE-US-00002 TABLE 1 Comparison of the results from the example calculations. Distance of Electricity Pyrolytic Investment-relevant Carbon foot- ws-SMR from Natural gas kWh .sub.el/kg carbon heat-transfer print H2 filling station kg CH4/kg H2 H2 kg C/kg H2 kWh .sub.th/kg H2 kg CO2/kg H2 Prior Electrolysis 50.0 24.4 9.5 art Mini-SMR  0 km 3.1 0.2 27.4 8.5 ws-SMR + H2 200 km 3.2 0.6 27.4 8.7 transport 500 km 3.2 0.6 27.4 9.2 Inventive ATP & EHS 3.5 6.0 1.1 10.0 6.7 examples LT plasma & 4.0 10.4 3.0 9.4 2.0 EHS RH pyrolysis & 4.0 8.4 3.0 19.2 1.6 EHS

REFERENCES

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