ES2737875A1 - CO2 CAPTURE INSTALLATION AND PROCEDURE (Machine-translation by Google Translate, not legally binding) - Google Patents

Abstract

Installation and procedure for capturing CO2 . The installation comprises a separation unit (1) by membranes into which a first air stream (2) enters which exits an oxygen enriched air stream (3) with an oxygen concentration of between 25-50% v/v . It comprises a boiler (4) to which the oxygen enriched air stream (3) and the fossil fuel (5) to be burned and from which a stream of flue gases (6) with between 25 -60% v/v of CO2 that is directed to a carbonator (8) of a calcination-carbonation unit (7) which also comprises a calciner (10) from which a CO current flows2 of high purity (11). A procedure for capturing CO2 in said installation is also described. (Machine-translation by Google Translate, not legally binding)

Description

D E S C R I P C I Ó N

CO2 CAPTURE INSTALLATION AND PROCEDURE

OBJECT OF THE INVENTION

The present invention is part of the technical field of technologies and processes oriented to the capture and storage of CO2 (CCS, carbonone capture and storage) released by the combustion of fossil fuels, biomass and its derivatives in industrial processes.

The installation can be used in plants that work all the processes destined to the production of electrical energy (power plants), cement, chemical industry or industrial heating systems, including district heating. The proposed installation and procedure integrate processes of membrane separation, partial oxy-combustion and CO2 capture after combustion through a calcination-carbonation cycle ( Calcium-Looping).

BACKGROUND OF THE INVENTION

The development of new efficient and low-cost CO2 capture and storage (CCS) technologies is a fundamental objective in order to mitigate the effects of climate change. The need to implement such technologies on a commercial scale is considered an urgent measure to limit global warming to 2 ° C above pre-industrial levels.

The projections made by the International Renewable Energy Agency (IRENA) contemplate a great penetration of renewable energies in the short term. However, the current and future energy system in the short-medium term still requires the use of fossil fuels in order to meet the enormous growth in the demand for energy that takes place worldwide. In addition, the increase in demand in emerging countries is especially notable.

CO2 capture technologies are expected to be vital for the decarbonization of the power production industry, with an estimate of 52 Gt captured from CO2 from 2015 to 2040. Plants equipped with CCS would generate benefits of 1.3 trillion dollars for coal and gas producers in that period of time (from 2015 to 2040).

CO2 capture technologies can be grouped into three broad categories: pre-combustion, post-combustion and oxy-combustion.

Pre-combustion CO2 capture is based on the reaction of the fuel with oxygen and / or air to produce synthesis gas (H2 and CO). Said reaction is carried out either in an atmosphere with a reduced content of O2 (partial or incomplete combustion) or in an atmosphere of water vapor. Subsequently, the CO reacts with water vapor in a catalytic reactor to produce CO2 and H2 through a reaction called Water Gas Shift (WGS). The CO2 can be separated from the H2 stream by physical or chemical means to achieve a H2 rich fuel stream.

The advantage of pre-combustion systems over other capture systems is that the CO2 / H2 separation process can be carried out using mature technologies and the resulting CO2 is at high pressure (15-40 bar) which reduces the compression work for storage As a main disadvantage is a high energy penalty of the process, which can be higher than 15 percentage points.

The capture of CO2 after combustion allows to obtain percentages of capture greater than 90% with a lower energy penalty than pre-combustion systems. These systems are based on separating CO2 from a flue gas stream, which has about 15% v / v CO2 in the environment if coal is used as fuel (coal is the most polluting of fossil fuels).

Among the main post-combustion CO2 capture techniques, the absorption processes using amines, Calcium-Looping and membrane separation can be highlighted. Among them, amine absorption technology is the only one currently tested on an industrial scale. Although percentages of CO2 capture can be obtained around 90%, absorption by amines has several drawbacks that limit its commercial development, such as a high energy penalty (8-12%) due to the regeneration of the sorbent as well as the degradation and the toxicity of amines. Thus, the specific energy consumption (called SPECCA) of amines systems would be of the order of 4 MJ / kgCO2.

The calcination-carbonation cycle ( Calcium-Looping ) is considered one of the main options within a new generation of advanced technologies for the capture of CO2 after combustion. Calcium-Looping is based on the reaction of CaO particles with the CO2 present in the flue gas stream through the carbonation reaction. One of the main advantages of this system is its low cost, as well as the wide availability and non-toxicity of natural CaO precursors, such as limestone (mostly composed of CaCO3) and dolomite. The use of other natural CaO precursors is also contemplated, such as steel slag, Ca-based synthetic precursors or compounds comprising inert materials resistant to sintering.

Once the carbonation reaction has elapsed, the generated CaCO3 is taken to a new reactor where the CaO is regenerated through a calcination reaction for a new cycle. The gases at the exit of the carbonation reactor (with approximately 90% less CO2) are discharged into the atmosphere while the gas released in the regeneration stage, mainly composed of CO2 and water vapor, is purified and compressed for storage or later use. Another great advantage of Calcium-Looping is that it has a lower energy penalty (6-8%) than other capture technologies.

The separation of CO2 from the post-combustion gas by membranes is undergoing extensive research. However, on the one hand, higher levels of permeability are necessary to treat large amounts of gases such as those produced in large thermal power plants. On the other hand, a cost reduction is necessary for the commercial development of the membranes as a post-combustion capture technique.

The capture of CO2 by oxy-combustion systems is based on carrying out combustion in a pure atmosphere (> 95% v / v) of O2. As a result, the flow of combustion gases is mainly composed of CO2 (> 70% v / v) and water vapor. After a purification process, whereby water vapor is mainly separated by condensation, the gas flow with a CO2 concentration of around 95% can be stored or used in other applications. This technology has been successfully tested on an industrial scale (30 MWth).

The main disadvantage of the oxy-combustion system is again the energy penalty that the capture process imposes on the performance of the plant, being in the range of 7-13 percentage points. The main cause of this high energy penalty is the energy consumption required for the production of very high purity O2 in the air separation units ( ASU). The two main techniques for the production of pure O2 are cryogenic distillation and the technology called pressure-swing adsoprtion ( PSA). The first one (cryogenic distillation) is commercially developed, with its estimated energy consumption for the production of O2 of around 200kWh / tn O2. On the other hand, PSA technology is developed for a smaller scale than cryogenic distillation with an also high energy penalty that prevents its development for CO2 capture.

Some hybrid processes for CO2 capture have also been proposed, such as the integration of oxy-combustion and absorption by amines or the integration of oxy-combustion with membranes for post-combustion capture.

For the commercial development of CO2 capture technologies it is essential to reduce the energy consumption associated with the capture process, that is, the specific energy consumption ( SPECCA) of the technology. On the other hand, the economic cost of capture systems should be as small as possible, for which a reduction in the sizes of current systems is necessary.

DESCRIPTION OF THE INVENTION

The present invention describes a CO2 capture facility in industrial processes and / or in electric power production plants (power plants) based on combustion systems.

The installation allows efficient integration of membrane, partial oxy-combustion and calcination-carbonation units. It is an integrated installation in which each of these units is combined with the others in synergy to reduce the overall energy penalty of the CO2 capture procedure performed at the installation, as well as to reduce the size of the system capture.

The installation comprises a membrane separation unit to produce an oxygen enriched air stream ( OEA) with an O2 concentration between 25-50% v / v.

The process of enrichment of O2 through a membrane system is economically interesting when moderate final concentrations of O2 (between 25-50%) are desired. As previously described, in order to produce greater purity in the O2 stream, for example when operating under conditions of total oxycombustion (> 95% v / v O2), the technically and economically viable options are cryogenic distillation or PSA that They have a high penalty of energy consumption.

Membrane systems for oxygen separation have great potential and industrial interest due to lower energy consumption and lower cost than cryogenic distillation plants. These systems are in commercial stage in small-scale plants (<50 ton O2 / day) while for larger plants they are in development phase. Due to the small difference in size between the oxygen and nitrogen molecules (3.46 Á and 3.64 Á, respectively) the process is technically complicated considering a simple separation based on diffusion across the membrane if a high purity of the generated O2 is required.

The most developed membranes currently for the separation of gases are polymeric membranes. These membranes have two advantageous characteristics to be applied in industrial-scale air separation processes in which high gas flow rates are required. The first of these is its ability to be prepared in hollow fiber modules, which gives the membrane a high packing density (up to 30,000 m2 / m3). The other is its modular nature, which allows its scaling without major modifications. According to previous studies, polymeric membrane systems for air separation are very efficient if high selectivity is not required, and the process can be configured in a single stage of membrane passage to produce enriched air with 40% v / v O2 .

Preferably, the membrane separation unit of the installation comprises polymeric membranes and is based on the diffusion mechanism across the membrane by applying a pressure difference that acts as the driving force of the process. Gas separation membranes are usually characterized by their permeability values (flow of a gas that crosses the membrane per unit of time) and selectivity, which represents the ability of a membrane to separate two gases.

The permeability of the gas (Pe¿) through the membrane can be calculated according to the following equation:

Nd

Pet = ----- l—

P2 - P 1

Where:

Nt = gas flow rate that passes through the membrane

1 = membrane thickness

Pi,> Pi = pressures of the gas before and after its passage through the membrane

In the case of polymeric membranes, as permeability increases, selectivity decreases, which implies reaching a compromise between both characteristics to achieve optimum values. Currently the highest values of selectivity factors for air separation with polymeric membranes range between 6-8.

The installation also includes a boiler into which the oxygen enriched air stream ( OEA) enters with an O2 concentration between 25-50% v / v leaving the separation unit by membranes. In the boiler the fuel is burned in partial oxycombustion mode. At the outlet of the boiler there is a stream of flue gases enriched in CO2 with respect to the case of combustion in air.

As a result of the partial oxy-combustion carried out in the boiler, a certain concentration of CO2 in the flue gases is obtained, depending on the concentration of oxygen in the oxygen enriched air stream and the fuel used. This concentration can be relatively high compared to that obtained by combustion with air. Partial oxy-combustion has been studied in the state of the art to improve combustion efficiency in boilers or to improve coal gasification performance. In the process of partial oxy-combustion of the fuel, heat is extracted which is used by means of a power cycle for the production of electricity. As in the case of oxy-combustion Total gas recirculation is necessary to control the flame temperature in the boiler.

The installation also includes a calcination-carbonation unit to which the flue gas stream is sent from the boiler. The calcination-carbonation unit comprises a carbonator and a calciner.

In an exemplary embodiment, if the installation is configured to have a CO2 concentration at the inlet of the calcination-carbonation unit of 30% v / v, the O2 concentration at the outlet of the membrane separation unit should be of ~ 38% v / v.

The flue gas stream is first sent to the carbonator, which is a fluidized bed reactor in which the flue gases enter at a rate of the order of 3-5 m / s. In the carbonator, the carbonation reaction occurs between the CO2 present in the flue gas stream and the solid particles of calcium oxide (CaO).

The endothermic carbonation reaction is carried out at atmospheric pressure and at a temperature between 625-680 ° C:

As a result of carbonation, calcium carbonate is produced. The carbonation however does not take place completely since a part of the calcium oxide remains unreacted. Thus, the conversion of CaO is defined as the mass ratio converted into CaCO3 between the total sorbent mass (CaO) before carbonation.

According to previous studies carried out by the authors of the present invention, carbonation, and therefore the CO2 capture efficiency is greatly favored by the increase in the concentration of CO2 in the flue gases. This improvement occurs mostly in the stage of carbonation limited by the kinetics of the reaction at the surface of the solids, which favors the operation of the process with high flows of solids that favor rapid conversions.

The calcination-carbonation unit also comprises a calciner where the carbonator outlet solids are sent. The calciner can also be a fluidized bed reactor or a flash calciner . Through a calcination reaction the sorbent (CaO) is regenerated from CaCO3 to be used in a new cycle.

As a result of the integration of the separation unit by membranes, the boiler for partial oxy-combustion and the calcination-carbonation unit, the installation proposed in the present invention has a specific energy consumption ( SPECCA) around 3 MJ / kgCO2, This indicates a huge potential for such integration as CO2 capture technology.

Taking into account the annual emissions of a typical 500 MWe power plant based on the combustion of coal, the energy savings when implementing the described installation compared to an installation based on CO2 capture with amines (SPECCA = 4 MJ / kgCO2) It could be around 1100 GWh / year. In addition, the installation has a capture efficiency of 95%, which allows to design reactors and smaller systems due to the reduction of the inventory of solids necessary to reach capture levels above 95%. This directly leads to savings in the cost of installation. Thus, the installation described represents a significant improvement in the CO2 capture process.

In the separation unit by membranes there is an energy consumption that is necessary to cause the pressure difference on both sides of the membrane. A first option to achieve said pressure difference is to increase the inlet pressure to the separation unit (before the membrane) with the area downstream of the membrane at a lower pressure (for example at atmospheric pressure). A second option to achieve said pressure difference is to consider atmospheric pressure before the membrane and impose vacuum downstream of the membrane.

The selection of the first or the second option depends on the configuration of the separation unit itself and the needs of the process carried out in the installation between which the input flows, temperatures, the CO2 concentration that can be varied can vary you want to obtain at the exit of the installation, selectivity of the membrane, cut-off factor (defined as the ratio between the amount of gas that passes the membrane between the total amount of gas at the entrance) of the separation process, etc.

According to the above, a study should be carried out before selecting one of the two options described for use in the installation. A priori, the strategy of imposing a partial vacuum in the area after the membrane would lead to lower energy consumption since the volume of gas in this area (the gas that passes through the membrane) is smaller than the gas entering the system.

In the separation process that is carried out in the separation unit using membranes and selecting as a separation strategy the creation of vacuum in the area enriched in O2 (after the membrane), the energy consumption of the process can be estimated from the calculation of the adiabatic work of ideal gas compression. In this way, in the separation process within the proposed installation, compression work is considered to impose the pressure difference on both sides of the membrane (at atmospheric pressure before the membrane, in vacuum on the side of the flow of oxygen enriched air) and the compression work to bring the gas from the oxygen enriched stream (output current of the separation unit) back to suitable boiler inlet conditions. These conditions require that the pressure be slightly higher than atmospheric to overcome the loss of load produced in the boiler.

Thus, the energy consumption of the separation process of the separation unit by membranes ( E sep) is calculated by the following expression:

Where:

mp = mass of gas (in this case oxygen) that crosses the membrane

Y = gas adiabatic expansion factor

T = is the system temperature

P¡n, Pp, Pb = pressures at the entrance of the membrane, at the exit of the membrane and at the entrance of the partial oxy-combustion boiler, respectively.

Any polymeric membrane can be used for the air separation process, although the efficiency of the process will be greater with membranes whose O2 / N2 selectivity is at least 4 with a permeability of at least 250 Barrers. Some examples of Polymeric membranes that fulfill these characteristics, and that therefore could be used are the intrinsic microporosity membranes (PIM membranes).

As previously described, the oxygen enriched air stream (25-50% v / V O2) is sent to the boiler to perform a partial oxy-combustion of coal or other fossil fuel (such as natural gas or biomass). Oxy-combustion generates the heat necessary for the production of electrical energy in a similar way to what would be done by a combustion process typical of a coal-fired power plant as fuel.

The boiler where oxy-combustion is carried out comprises a set of heat exchangers for steam production at certain conditions of live steam required in a turbine of a primary power unit, which is also part of the installation. This set of heat exchangers is of the steam generator type (HRSG). At the exit of this set of heat exchangers the combustion gases have a temperature of around 200 ° C.

The combustion gases at the outlet of the boiler are directed to the carbonation calcination unit ( Calcium-Looping). The installation may comprise a compressor at the entrance of said calcination-carbonation unit to overcome the loss of system load. It can also comprise a heat exchanger, gas-gas type, at the inlet of the calcination-carbonation unit to transfer heat between the output gases of said unit (these output gases are poured directly into the atmosphere) and the gases of entrance of said unit (the combustion gases that come from the boiler).

When the gases to be introduced into the calcination-carbonation unit are at the appropriate temperature, they are passed to the carbonator which, as previously described, is a fluidized bed reactor. Here the carbonation reaction occurs between the CO2 present in the flue gases and CaO particles. To maximize the transfer of heat and mass, the carbonator can operate in a circulating fluidized bed regime, with superficial gas velocities around 3-5 m / s.

One of the main disadvantages of the Calcium-Looping cycle is the progressive loss of sorbent (CaO) activity in the carbonator with the passing of the cycles. This is mainly due to the sintering of the sorbent, which causes the reduction of its specific surface for the solid-gas reaction in short residence times.

In the carbonation reaction two different stages can be found: a first stage dominated by the carbonation kinetics that ends with the formation of a CaCO3 layer that covers the surface of the particles of approximate thickness around 40-60 nm and a second stage, whose reaction rate is noticeably slower, in which the CO2 penetrates through the carbonate layer by means of a diffusion mechanism. As a consequence of the usual residence times in the industry for this type of reactors (between 2 and 10 minutes), the nucleus of the particles remains unreacted, thus leaving part of the bed as CaO without reacting with the consequent loss of system performance . In operation on an industrial scale, solids are cycled hundreds of times, their conversion into carbonation falling to a residual value of the environment of Xr = 0.06.

The carbonation reaction in the Calcium-Looping process of the calcination-carbonation unit typically takes place at temperatures around 650 ° C, which favors the conversion kinetics at atmospheric pressure while at the same time the concentration of CO2 in the equilibrium sufficiently low (~ 1% v / v).

In the proposed installation, the combustion gases enter the carbonator with an increased CO2 concentration, so carbonation is favored. Thanks to this, greater capture efficiency is obtained. Specifically, the kinetic phase of carbonation is highly potentiated, which allows for greater conversion, while conversion in the diffusive phase is slightly penalized. Thus, it is observed that increasing the concentration of CO2 in the flue gases (by means of the partial oxy-combustion that has been carried out) achieves a greater residual conversion of CaO in the carbonation (Xr = 0.08 in the case of having a concentration of CO2 at the entrance of the calcination-carbonation unit of 60% v / v), and therefore a greater capture of CO2 (up to 95%).

On the other hand, it is observed how increasing the carbonation temperature up to 680 ° C achieves a higher conversion (for example, 15% more in the case of having a CO2 concentration at the entrance of the carbonation calcination unit of 30% v / v,) in carbonation thanks to an increase in the extension of the reaction in the diffusive phase (the behavior of the kinetic phase is similar). This also occurs in the calcination-carbonation unit when the CO2 concentration is around 15% v / v, in the latter case being the most important effect (residual conversion varies by 21%). This is an important improvement since variations in the temperature in the range of 625-680 ° C can occur, as a result of specific defects in heat transfer, they would penalize less in the case of the present installation because the effect of The carbonation temperature has less influence.

The heat generated in the carbonation (exothermic reaction) is used for the production of superheated steam that feeds a secondary power cycle (or the same power cycle in which the previously described turbine is located) in which it is carried out a Rankine cycle for the production of electric power. Thus, the installation can have up to two power production units, a primary power unit that uses as energy source the energy released in partial oxycombustion and a secondary power unit that uses the heat released in carbonation.

At the exit of the carbonator, the combustion gases have up to 95% less CO2 and a temperature of around 650 ° C. Before being discharged into the atmosphere, the gases pass through a second gas-gas heat exchanger to preheat the incoming water stream to the secondary steam generator and to the flue gas stream entering the carbonator.

At the exit of the carbonator there is also a stream of solids (separated from the gaseous stream by means of a cyclone located at the outlet of the carbonator). This stream of solids is sent to the calciner where a regeneration of the sorbent is produced by the calcination reaction (reverse endothermic reaction to that of carbonation).

Prior to the inlet of the calciner, the stream of solids leaving the carbonator is passed through a second heat exchanger to increase its temperature taking advantage of the high temperature of the solids leaving the calciner. This solid-solid heat exchanger can be configured as two solid-gas exchangers in series. By means of this exchanger, the outgoing solids of the carbonator (CaCO3 and CaO that has not reacted) are heated up to a temperature around 900 ° C. In an exemplary embodiment, given the loss of solids activity, "fresh" limestone can be added to the system (~ 1-2%) with the consequent purging of cycled material to maintain the mass balance.

In the calciner, a high CO2 concentration and therefore a temperature around 950 ° C are established as atmospheric pressure conditions to ensure that calcination occurs entirely in short residence times. Although the temperature of calcination at atmospheric pressure according to the thermodynamic equilibrium is 896 ° C under an atmosphere of pure CO2, it has been experimentally proven that for the reaction to be carried out with a kinetics appropriate to the residence times of the reactor , the temperature must be higher than 930 ° C, considering an atmosphere of pure CO2. The heat necessary to carry out the calcination, and to increase the temperature of the solids entering the calciner to the calcination temperature, is obtained from a process of total oxy-combustion of carbon in-situ in the calciner.

In an exemplary embodiment in which a cryogenic distillation process for the production of O2 with high purity (> 95%) is carried out, this production of O2 causes a significant penalty to the process (~ 25% of the total power consumption of Calcium -Looping). However, it is necessary to generate high purity O2 for the total oxy-combustion process, from which a gas stream with a CO2 concentration of ~ 75-80% v / v is obtained at the outlet of the calciner, which is it passes through a purification unit (basically consisting of the separation of water vapor by condensation) to increase its CO2 concentration to ~ 90-95% v / v. Thus the CO2 stream is ready for storage or later use in another process. In the case of storage, the CO2 must be compressed up to ~ 100 bar for example by a compression process with intermediate cooling.

The main advantage of the present invention is that a lower energy penalty is achieved than the currently known CO2 capture processes. There is also a lower consumption of specific energy used in CO2 capture compared to CO2 capture processes of the prior art. As previously seen, with the described installation, consumption of 3 MJ / kg CO2 can be achieved compared to consumption values around 3.5 MJ / kg CO2 or 4 MJ / kg required for example when Calcium-Looping and oxy - combustion processes are used independently.

Furthermore, in the proposed installation, carbonation is more stable against temperature changes in the 625-680 ° C range than in the case of Calcium-Looping performed independently. Also, in comparison with said procedure, a higher capture efficiency is obtained, which implies a smaller amount of CaO in the bed of the calciner, with the consequent decrease in the size and cost of the reactors.

The described installation is scalable from 1 MWe to more than 500MWe.

DESCRIPTION OF THE DRAWINGS

To complement the description that is being made and in order to help a better understanding of the characteristics of the invention, according to a preferred example of practical implementation thereof, a set of drawings is attached as an integral part of said description. where, for illustrative and non-limiting purposes, the following has been represented:

Figure 1.- Shows a conceptual scheme of the CO2 capture facility.

Figure 2.- Shows a detailed scheme of the installation.

PREFERRED EMBODIMENT OF THE INVENTION

Examples of embodiments of the present invention are described below with the aid of Figures 1 to 2.

Figure 1 shows a diagram of the units that are part of the installation and the flows that move between them.

As can be seen in said figure, the installation comprises a separation unit (1) by membranes for the production of a stream of oxygen enriched air into which a first air stream (2) enters enriched in oxygen (3). with an oxygen concentration of between 25-50% v / v.

The installation also has a boiler (4) for partial oxy-combustion of a fossil fuel by means of the oxygen enriched air stream obtained in the separation unit into which the oxygen enriched air stream (3) and the fossil fuel ( 5) that is going to burn and from which a stream of flue gases (6) comes out with between 25-60% v / v of CO2.

Another essential unit of the installation is a calcination-carbonation unit (7) for the capture of CO2 from the flue gas stream (6) by means of a Calcium-Looping process and comprising a carbonator (8) to which the flue gas stream (6) and from which a "clean" gas stream (with a significantly lower amount of CO2) (9) and a stream of solids (separated from the purified gas stream (9) preferably with a cyclone), and comprises a calciner (10) which reaches the stream of solids that comes out of the carbonator (8), together with fuel and oxygen, and from which a stream of pure CO2 (11) leaves for compression, storage or reuse The thermal energy necessary to carry out the calcination is obtained by total oxy-combustion of coal, biomass or other similar fuel.

As an example of the invention, a CO2 capture process is shown by the proposed installation in a 500MWe coal thermal power plant. Figure 1 shows a conceptual scheme of the installation to facilitate the understanding of the procedure. A detailed diagram of the installation is shown in Figure 2. The most relevant data regarding matter and energy balances in the system are shown in Table 1.

In this example the air stream enriched in O2 at the outlet of the separation unit by membranes (1) is 281 kg / s with an O2 concentration of 38% v / v. For this, it is necessary to process 2810 kg / s of air considering a cut-off factor of 0.1.

In the installation, an energy consumption is assumed in the process caused by the vacuum pump and the subsequent compression of the enriched air stream in O2 (3) up to the boiler inlet conditions (4) of 90 kWh / tn O2, equivalent consumption to use a vacuum pressure of 0.25 bar.

The oxygen enriched air stream (3) enters the boiler (4) where the partial oxy-combustion of the fossil fuel (5) occurs, which in this embodiment is 42.2 kg / s of Pittsburg No. 8 coal. At the outlet of the boiler (4), the flue gas stream (6) has a CO2 concentration of 30% v / v. The efficiency of the boiler (4) assumed is 88%, and the mass flow of O2 is such that a fraction of O2 (excess of combustion oxygen) of 2.5% is obtained in the flue gas stream (6) v / v.

To control that the flame temperature inside the boiler is below 1400 ° C, a recirculation of combustion gases of 63% is imposed. A compressor through which the recirculation current passes ensures the necessary pressure at the boiler inlet. In a preferred embodiment, a first cyclone is available at the outlet of the boiler to separate ashes and possible unburned solids from the gas stream.

The heat released in the combustion carried out in the boiler (4) is used in an HRSG type heat exchanger (12) to produce superheated steam in a primary steam cycle (24), where a preheated water stream (13) (168 ° C, 300 bar) evaporates and overheats until the imposed live steam conditions (600 ° C, 290 bar) are reached. The primary steam cycle (24) is configured with an intermediate reheating stage in a third exchanger (20) in which a steam stream (21) passes from 188 ° C to 620 ° C, this heat also being obtained by the boiler (4) of oxy-combustion. The superheated steam produced by the heat released in the boiler is used to obtain energy (26) in a turbine that is part of the first main power unit (14).

The condensation pressure of the primary steam cycle (24) is 0.05 bar. As a consequence of the above, the primary steam cycle (24) has a thermoelectric efficiency of 45%. The net power resulting from the oxy-combustion process is 513 MWe, with a complete system yield of 36.2%. Taking into account the performance of the main power unit (14) in the case of combustion with traditional air in coal-fired power plants, the penalization of the process carried out in the separation unit by means of membranes (1) and the partial oxy-combustion in the boiler ( 4) is 1.63%.

At the outlet of the oxy-combustion boiler (4) a part of the gas stream is recirculated (28) to keep the flame temperature in the boiler below 1400 ° C while 322.9 kg / s of flue gas (6 ) with a CO2 concentration of 30% v / v are sent to the calcination-carbonation cycle. The installation comprises a cooling system to reduce the amount of water vapor in the flue gas stream (6) and to reduce the energy consumption in a first compressor (22) arranged at the inlet of the carbonator (8) that ensures that said flue gas stream (6) reaches the carbonator (8) with the required pressure conditions.

To increase the temperature of the flue gas stream (6) entering the carbonator (8), there are two gas-gas heat exchangers (24, 25) in series that, for heating, use respectively: sensible heat from the stream of pure CO2 (11), compressed and ready for storage after the Calcium-Looping process (at the outlet of the calciner (10)) and the heat of the stream of purified gas (9) at the outlet of the carbonator (8) ready to be poured into the atmosphere. Thus, the entry of gases into the carbonator (8) occurs at a temperature of 198 ° C and with a pressure of 1.2 bar, sufficient to cover the pressure losses that occur both in the carbonation reactor and in the cyclone at the carbonator outlet (the cyclone in which the stream of solids is separated from the gaseous stream leaving the carbonator).

In the carbonator (8), the flue gas stream reacts with an incoming flow of solids (18) of 2535 kg / s (2388 kg / s of CaO while the rest are ashes and CaSO4 formed in the process) at 670 ° C from the calciner (11). This assumes a molar flux ratio of solids versus molar flux of CO2 FR / F C02 = 15.

Considering an inventory of solids in the calciner (8) of 400 tons, the capture efficiency obtained is 95%. Thus, an amount of CaCO3 of 270 kg / s is produced while 2236 k / s of CaO remain unreacted. The purified gas stream (9) is composed of flue gases at the outlet of the carbonator (8) that have a molar fraction of 0.02 CO2, which makes its discharge into the atmosphere acceptable after taking advantage of some of its sensible heat to preheat the water stream (16) to the input of the secondary power unit (17) (steam generator) up to 329 ° C in the second heat exchangers (15) and for preheating the inlet gases to the carbonator (8), which are the flue gases (6) protruding from the boiler (4), in the first heat exchanger heat (12).

The heat produced in the exothermic carbonation reaction is used to feed a secondary steam cycle (second power unit (17)), obtained 177 kg / s of superheated steam with live steam conditions of 600 ° C and 290 bar. In the secondary steam cycle (25) the reheating stage is carried out using part of the sensible heat of the pure CO2 stream (11) at 950 ° C leaving the calciner (10). With the above, from the secondary steam cycle (25) an energy (26) corresponding to a net power of 288 MWe is obtained. An important part of this electrical power (~ 60%) is used in the calcination-carbonation unit (7), taking place the main consumption in the production of O2 necessary for the partial oxycombustion of the calciner (8), in the transport of solid and in the compression of the captured CO2 for storage or use in other industrial processes.

After being separated from the gas stream, the solids (18) are sent to the calciner (8) for the regeneration of the sorbent, previously passing through a heat exchange system (26) that causes the temperature rise of the solids inlet ( 18) up to 912 ° C at the inlet of the calciner.

This significantly reduces the fuel consumption intended to increase the temperature of the reagents to the conditions necessary for calcination (950 ° C and atmospheric pressure). In the calciner (8), the oxy-combustion of 22.3 kg / s of fuel (coal) is necessary to carry out the heating of the CaCO3 so that the endothermic calcination reaction takes place, by which the sorbent (CaO) is regenerated. It will then be sent to the carbonator (8) for a new cycle. In other embodiments the fuel may be different from coal and the CO2 source may be another type of fossil or industrial fuel plant.

In the calciner (8), a current of 59.5 kg / s of O2 (95% v / v) produced by a cryogenic distillation process is used for oxy-combustion, with a consumption associated with the 200 kWh / tn O2 process produced.

At the outlet of the calciner (10) a gas stream is obtained, which has been referred to as pure CO2 stream (11), of 203.4 kg / s compound mainly of CO2 (90% v / v) and water (5% v / v). After assigning part of its sensible heat in the reheating of the secondary steam cycle by means of a superheater (27) it is passed through a purification system to remove, by condensation, the water vapor while the CO2 is compressed in a compressor ( 23) up to 100 bar for storage or later use.

As a consequence of the process carried out in the carbonation calcination unit (7), an energy penalty of 6.12% is obtained. This penalty, together with the 1.63% penalty caused by the separation unit using membranes (1) in the stage prior to partial oxy-combustion in the boiler (4), results in a specific fuel consumption destined for the capture system in the complete process that is carried out in the installation of 3.05 MJ / kg CO2, lower value than that obtained in CO2 capture processes by oxy-combustion (~ 4 MJ / kg CO2) or Calcium-Looping (~ 3.5 MJ / kg CO2).

The following table shows the values of different parameters obtained using a CO2 capture facility as described.

Table 1: Main simulation results of the CO2 capture facility

Where:

Claims (16)

1. - CO2 capture facility characterized by comprising: - a separation unit (1) by membranes, which comprises at least one membrane, for the production of an oxygen enriched air stream to which a first air stream (2) enters which an oxygen enriched air stream exits ( 3) with an oxygen concentration of between 25-50% v / v; -a boiler (4) for partial oxy-combustion of a fossil fuel by means of the oxygen enriched air stream obtained in the separation unit into which the oxygen enriched air stream (3) and the fossil fuel (5) enter will burn and from which a stream of flue gases (6) with 25-60% v / v of CO2 comes out, -a calcination-carbonation unit (7) for the capture of CO2 from the flue gas stream (6) by means of a Calcium-Looping procedure and comprising a carbonator (8) to which the flue gas stream enters ( 6) and a stream of solids rich in CaO from the calciner (18) and from which a stream of purified gas (9) and a stream of solids rich in CaCO3 comes out, and comprises a calciner (10) to which the stream of solids that leave the carbonator (8) and a stream of O2 to perform the calcination by oxy-combustion of a fuel, and from which an almost pure CO2 stream (11) comes out for compression, storage or reuse. 2. - CO2 capture installation according to claim 1 characterized in that at least one membrane of the membrane separation unit (1) is a polymeric membrane. 3. - Installation of CO2 capture according to claim 2 characterized in that the membranes have an O2 / N2 selectivity of at least 4 and a permeability of at least 250 Barrers. 4. - Installation of CO2 capture according to claim 1 characterized in that the solid fuel (5) entering the boiler is selected from carbon, natural gas or biomass. 5. - Installation of CO2 capture according to claim 1 characterized in that the flue gas stream (6) has a CO2 concentration greater than 15% v / v. 6. - Installation of CO2 capture according to claim 1 characterized in that the carbonator, the calciner or both are fluidized bed reactors. 7. - CO2 capture installation according to claim 1 characterized in that it comprises a vacuum pump between the separation unit (1) and the boiler (4) configured to create a vacuum that forces the passage of the first air stream ( 2) through the membrane. 8. - Installation of CO2 capture according to claim 1 characterized in that it comprises a plurality of first heat exchangers (12) through which the flue gas stream (6) passes to heat a first stream of water vapor (13) to be used in a turbine in a main power unit (14). 9. - Installation of CO2 capture according to claim 1 characterized in that it comprises a compressor at the inlet of the calcination-carbonation unit (7) through which the flue gas stream (6) passes before entering the calcination-carbonation unit (7). 10. - Installation of CO2 capture according to claim 1 characterized in that it comprises a second heat exchanger (15) arranged in the carbonator to heat a second stream of water vapor (16) to be used in a turbine in a unit of secondary power (17). 11. - CO2 capture installation according to claim 1 characterized in that it comprises a first cyclone arranged at the outlet of the boiler (4) to separate ashes and unburned solids from the flue gas stream (6). 12. - Installation of CO2 capture according to claim 1 characterized in that the sorbent in the carbonator (8) is CO. 13. - CO2 capture procedure in a CO2 capture facility as described in one of claims 1 to 12, characterized in that it comprises the following steps: a) introducing the first air stream (2) into the membrane separation unit (1); b) apply a pressure difference between the two sides of the membrane to force the passage of the first air stream (2) through said membrane; c) introduce the oxygen enriched air stream (3) and the fossil fuel (5) into the boiler (4) and perform a partial oxy-combustion of the fossil fuel (5); d) introduce the flue gas stream (6) into the carbonator (8) and a stream of solids rich in CaO; e) introduce a stream of solids from the carbonator rich in CaCO3 (8) in the calciner (10) and O2 to perform calcination by oxy-combustion and obtain a stream of high purity CO2 (11). 14. - Method according to claim 13 characterized in that it further comprises a step f) of condensing the water vapor of the high purity CO2 stream (11) and compressing said stream into a compressor (23). 15. - Method according to claim 13 characterized in that it comprises a stage of heating a first stream of water vapor (13) in the boiler (4) and a subsequent stage of using said first stream of water vapor (13) in the turbine of the main power unit (14) to obtain electric power. 16. - Method according to claim 13 characterized in that it comprises a step of heating a second stream of water vapor (16) in a second heat exchangers (15) of the carbonator (8) and a subsequent stage of using said second stream of water vapor (16) in the secondary power unit turbine (17) to obtain electric power.