Patent Application
20120301391
The present invention relates to a process for the production of hydrogen starting from liquid hydrocarbons, gaseous hydrocarbons, and/or oxygenated compounds, also deriving from biomasses, and mixtures thereof. Said process comprises:
Said process can possibly comprise a hydrodesulphuration section of said feedstock.
The most widely-used technology for the production of synthesis gas and subsequently of hydrogen is the Steam Reforming (SR) process. This technology transforms light desulphurated hydrocarbons, by reacting them with steam, in direct fired multitubular catalytic reactors, inserted in an oven, according to the reaction [1]:
CH4+H2O=CO+3H2ΔH°=49.3 kcal/mole [1]
The combustion serves to provide heat to the reactions which are extremely endothermic. The hydrocarbons enter the reforming tubes after being mixed with significant quantities of steam (the [steam moles/carbon moles] ratio is typically higher than 2.5) and are transformed into a mixture prevalently containing H2 and CO (synthesis gas). The catalysts used typically contain Nickel deposited on an oxide carrier. The inlet temperatures into the tubes are typically higher than 600° C., whereas the temperatures of the gases leaving the tubes are lower than 900° C. The pressure at which the SR process takes place typically ranges from 5 relative bar to 30 relative bar.
More specifically, the SR process takes place in a tubular reactor in which the tubes are inserted in a radiant chamber and in which the reaction heat is supplied through wall or vault burners. In the SR reactor, the reaction tubes have a diameter ranging from 3″ to 5″ and a length of 6 metres to 13 metres; said tubes are filled with catalyst and the mixture composed of hydrocarbons and steam passes through them.
In order to obtain the outlet temperatures of the synthesis gas within the range of [800-900]° C., the wall temperature of said tubes is about [100-150]° C. higher and that of the fumes generated by the burners is [1200-1300]° C. These tubes, constructed by fusion with special alloys having a high Cr and Ni content ([25-35]%), consequently represent a critical element of the technology. The necessity of avoiding impingement between the tubes and flames of the burners, which would lead to the instantaneous collapse of the tubes, requires their distancing and consequently an increase in the volume of the reforming oven. A further critical aspect of the SR process relates to the impossibility of using high-molecular-weight hydrocarbons, which can lead to the formation of carbonaceous residues with a reduction in the catalytic activity. As a result of this, the heat supplied to the outside of the tubes causes cracking phenomena of the hydrocarbons, with a further formation of carbonaceous residues, of which the most extreme consequence is the blockage of the reforming tubes and their breakage. The sulphurated compounds, if fed to the SR process, can also cause deactivation of the catalyst and create analogous consequences. For this reason, for the SR process, the feedstock must be hydro-desulphurated before being used.
From an operative point of view, in an environment such as a refinery, the management of an SR oven consequently creates a series of critical elements which are currently solved by a continual monitoring of the same.
Various configurations and technologies have been proposed for solving some of the critical aspects relating to the SR technology. One of these is represented by the short contact time—catalytic partial oxidation (SCT-CPO) process described in the patents MI93A001857, MI96A000690, MI2002A001133, MI2007A002209 and MI2007A002228 of L. Basini et al. In this technology, the hydrocarbons mixed with air and/or oxygen are passed over a suitable catalyst and transformed into synthesis gas. The reaction heat is generated inside the reactor, by balancing the total and partial oxidation reactions of the feedstock. When natural gas is used, the main reaction of the SCT-CPO process is represented by the equation [2]:
CH4+½O2=CO+2H2ΔH°=−8.5 kcal/mole [2]
This reactor is extremely simplified in its constructive and operative principles. The reactor is of the adiabatic type with dimensions over two orders of magnitude lower than the SR reactor. The catalysts, moreover, are not deactivated (unlike what takes place in the SR process) even if there are sulphurated compounds in the feedstock; this allows a process architecture in which the hydro-desulphuration step can be avoided. The constructive simplicity and resistance of the catalyst to deactivation phenomena also allow a considerable management simplicity and reduced maintenance interventions. More specifically, it is indicated that to produce 55,000 Nm3/hour of hydrogen with the SR technology, an oven containing 178 catalytic tubes is necessary. It is also estimated that, in this case, the volume of catalyst required amounts to about 21 Tons. It is also specified that the reaction section and thermal recovery section from the fumes of the reforming oven have considerable dimensions and occupy a volume of approximately 11,000 m3. The same quantity of H2 could, on the other hand, be produced by an SCT-CPO reactor and a thermal recovery section having a total volume of about 70 m3 and containing 0.85 Tons of catalyst.
In the SR process destined for the production of H2, the synthesis gas leaving the reforming oven is shifted to a mixture of H2 and CO2 by reacting the CO with water vapour in one or more Water Gas Shift (WGS) reactors according to the reaction [3]:
CO+H2O=CO2+H2ΔH°=−9.8 kcal/mole [3]
The H2 is subsequently separated and purified typically using a Pressure Swing Adsorption (PSA) section. The latter exploits the different physisorption properties of the molecules on different kinds of materials. The PSA section therefore releases a stream of pure H2 and a stream of low-pressure purge gas which mainly comprises CO2, CH4 and a part of the H2 produced. Said purge gas which has a heat power (PCI) typically within the range of [2,000-2,500] kcal/kg, it is then fed again to the reformer oven supplying a part of the reaction heat. One of the disadvantages of the SR reaction is the export production of steam, i.e. an excess production of steam which cannot be recovered in the process and whose presence reduces the energy efficiency of the process itself.
A similar process scheme can also be used in the SCT-CPO technology destined for the production of H2. In this case, however, the partial pressure of the CO2 produced at the outlet of the WGS section is higher than that obtained in the SR process, and consequently not only the flow-rate of the gas to be purified is higher in PSA, but also the purge gas leaving the PSA has a lower heat power with respect to that obtained by means of SR. A purge gas with an excessively low heat power value cannot easily be used for the production of steam in a boiler.
An objective of the present invention is to provide a new process architecture which combines a SCT-CPO section, a WGS section and a CO2 removal section in order to obtain a stream of H2, with purity higher than 90% v/v, separated from a stream of pure CO2. In a possible process configuration, in addition to the three previous sections, there is also a PSA section, situated after the CO2 removal section. This PSA unit allows high-purity, H2 and a purge gas with a medium heat power, to be obtained.
A further objective of the present invention is therefore to produce streams of high-purity H2 and CO2 and a purge gas leaving the PSA with a medium-high heat power (PCI), which is such as to allow it to be used directly, in combustion processes and/or introduced into the fuel supply system of a plant. Finally, specifically because the hydrodesulphuration step of the feedstock can be avoided, a further objective of the present invention is to allow the production of synthesis gas containing lower quantities of sulphurated compounds, which could be eliminated in the CO2 removal step and/or in the possible PSA step.
The present invention relates to a process for the production of hydrogen starting from reagents comprising liquid hydrocarbons, gaseous hydrocarbons, and/or oxygenated compounds, also deriving from biomasses, and mixtures thereof, wherein the gaseous hydrocarbons are selected from the group comprising natural gas, liquefied petroleum gas, gaseous hydrocarbon streams coming from operative processes in refineries and/or any chemical plant and mixtures thereof, wherein the liquid hydrocarbons are selected from the group comprising naphthas, gas oils, high-boiling gas oils, light cycle oils, heavy cycle oils, deasphalted oils, and mixtures thereof, and wherein the oxygenated compounds are selected from the group comprising glycerine, triglycerides, carbohydrates, methanol, ethanol, and mixtures thereof, said process characterized in that it comprises:
A further embodiment of the present invention relates to a process as previously described possibly comprising a purification section of the hydrogen produced by means of Pressure Swing Adsorption and the generation of purge gas having a medium heat power.
The purge gas can possibly be used in a combustion process and/or be introduced into the fuel supply system of a refinery or any other chemical plant. Having considerably reduced the flow-rate to the PSA, thanks to the removal of the CO2, the possible final purification of the hydrogen is more efficient and less costly. Furthermore, this process greatly reduces emissions such as NOx, CO and particulates, as the preheating of the feedstocks can preferably be effected with the steam produced by the cooling of the synthesis gas leaving the SCT-CPO reactor. Process schemes which adopt the synthesis gas production technology via SCT-CPO may also not use preheating ovens of the reagents; it is therefore always possible to avoid producing diluted streams of CO2 in the combustion fumes.
Finally, the process configuration can be such as to not cause the production of an excess of steam. The export of steam, in fact, is not always advantageous and in some cases it may be advisable to avoid it.
A further embodiment of the present invention relates to a process as previously described which possibly comprises a hydrodesulphuration section of the reagents.
The process integration between the hydrodesulphuration section, SCT-CPO, WGS reaction, CO2 removal and PSA can also be formulated so as to not cause any emission of CO2 in diluted streams different from that obtained from the removal unit. The SR technology, on the contrary, does not allow a process scheme to be formulated in which an overproduction of steam (we repeat that the export of steam in fact is not always advantageous or necessary in all industrial contexts) or the emission of CO2 in the fumes of the preheating and SR ovens, can be avoided. The quantity of CO2 emitted and “not recoverable” corresponds to percentages ranging from 30% v/v to 45% v/v of the total quantity of CO2 produced.
All of these advantages together make the production cost of hydrogen in different scenarios more competitive with respect to that which can be obtained with the conventional SR technology.
Further objectives and advantages of the present invention will appear more evident from the following description and enclosed drawings, provided for purely illustrative and non-limiting purposes.
205 and 206 obtain the production of steam to be exploited in the process.
According to what is represented in
With reference to
The feeding (2) comprises liquid hydrocarbons, gaseous hydrocarbons, and/or oxygenated compounds, also deriving from biomasses, and mixtures thereof. The gaseous hydrocarbons comprise natural gas, liquefied petroleum gas, gaseous hydrocarbon streams coming from operative processes in refineries and/or any chemical plant and mixtures thereof. The liquid hydrocarbons comprise naphthas, gas oils, high-boiling gas oils, light cycle oils, heavy cycle oils, deasphalted oils, and mixtures thereof.
The oxygenated compounds comprise glycerine, triglycerides, carbohydrates, methanol, ethanol and mixtures thereof.
The feeding (2) possibly enters the hydrodesulfphuration section (100) where the sulphur is initially converted to sulphidric acid and is subsequently reacted with zinc oxide so that the outgoing feedstock contains less than 0.1 ppm of sulphur. The hydrodesulfphuration section may not be the initial step of the process as the catalytic partial oxidation section (101) is capable of also operating with sulphurated feedstocks. The hydrodesulfphuration section (100) can be situated downstream of a Water Gas Shift Sulphur Tolerant section (not indicated in
More specifically, as far as the steam cycle is concerned, it has been observed that a part of the H.T. Steam (5), generated in the cooling of the stream of synthesis gas produced (4), is injected into the WGS section (102) to guarantee high conversions of the carbon monoxide and allow the formation of H2 and CO2 (9). The mixture obtained after the WGS reaction is cooled producing low-pressure steam (13 and 20), a part of which can preferably supply the heat necessary for the regeneration section of the amines possibly used in the CO2 removal section (104). In a further phase, the mixture of H2 and CO2 is cooled with water by means of a Boiling Feed Water cooler (202) and is then cooled with an air exchanger (203) and with a water exchanger (204) before being sent to a section which removes the condensate (103). After removing the condensates, the gas (9) is sent to the carbon dioxide removal section (104). The CO2 removal section preferably includes an amine washing section, but it can also include any other system. This section preferably removes at least 98% of the carbon dioxide contained in the synthesis gas. After the removal of the CO2, the gaseous stream obtained contains a high percentage of H2, preferably higher than 80% v/v, but even more preferably higher than 90% v/v, said stream can be treated by a PSA section having reduced dimensions (105). Said PSA section allows a high recovery factor of the H2 produced (16) to be obtained, higher than 85% v/v and preferably higher than 90% v/v. The total or almost total lack of CO2 in the stream which can be sent to the PSA significantly increases the heat power of the purge stream allowing it to be re-used in combustion processes and/or to be introduced into the fuel supply system of a refinery or any other chemical plant. In a preferred embodiment, part of the purge gas (3) is used as fuel for a preheating oven of the reagents (200), before entering the SCT-CPO section. The purge gas separated by means of PSA, in fact, has a relatively high heat power, with a value at least equal to 4,000 kcal/kg, preferably ranging from 4,500 kcal/kg to 7,000 kcal/kg and even more preferably ranging from 5,000 kcal/kg to 6,000 kcal/kg.
Table 1 compares the consumptions of two typical Steam Reforming and SCT-CPO plants, both structured for recovering CO2. The comparison is centred on the analysis effected for plants with a capacity of 55,000 Nm3/hour of H2. Example 1 refers to
1Calculated by subtracting the heat of the purge gas.
From a comparison between the total and specific consumptions, an extremely favourable situation emerges for the SCT-CPO technology if compared with the SR technology in the presence of CO2 recovery. More specifically, it can be noted that the consumptions of natural gas or rather the calories input per unit of product proves to be almost 4% lower for the SCT-CPO technology, with an emission of CO2 ten times lower, which leads this technology to be considered a winning choice when a CO2 recovery is to be installed. There are evident economical advantages which are even more so in contexts which jeopardize the production of CO2 and reward its “sequestration” and re-use.
It should be pointed out that in SR, an important part of the CO2, approximately a third, remains in the fumes and its recovery creates problems which are difficult to solve technically (degradation of the adsorbing solutions in the presence of oxygen) and which imply operative costs which are so high as to make this solution not to be proposable. In SR, a total recovery of the CO2 is consequently unconceivable as it can be done in the SCT-CPO where all the CO2 is present in the process gas.
The SCT-CPO technology, on the contrary, is jeopardized by a higher consumption of cooling water and electric consumption relating to the cryogenic unit for separating the air and obtaining pure oxygen. Between the two, the cost of electric energy is almost two orders of magnitude higher. The advantage of the SCT-CPO technology is consequently greater in countries in which the energy cost is lower. It should be noted that the advantage with respect to consumptions is additional to that relating to the investment costs, as the complexity of the synthesis gas production section is considerably reduced passing from the SR technology to the SCT-CPO technology.
In this example, reference is again made to
The specific consumptions were again evaluated using, for Steam Reforming, the data indicated by the licensees, and for the SCT-CPO technology, the consolidated data at a bench-scale level. The electric consumptions for the compression operations and separation of the oxygen in the Air Separation Unit are not included.
1Calculated by summing the natural gas at the burners.
As for Example 1, the process configuration adopted for the SCT-CPO process is clearly more advantageous in contexts in which the “sequestration” and re-use of CO2 is rewarding and in contexts in which the cost of electric energy is low.
Furthermore, in this case, the percentage reduction in the investment costs relating to the reduction in the complexity of the synthesis gas production section of the SCT-CPO process increases with respect to the SR process.
| Number | Date | Country | Kind |
|---|---|---|---|
| MI2009A 002199 | Dec 2009 | IT | national |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/EP2010/007772 | 12/15/2010 | WO | 00 | 8/3/2012 |
