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Integration of the Ca-Cu process in ammonia production plants Isabel Martínez, David Armaroli, Matteo Gazzani, and Matteo Carmelo Romano Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b04615 • Publication Date (Web): 03 Feb 2017 Downloaded from http://pubs.acs.org on February 18, 2017
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Industrial & Engineering Chemistry Research
Integration of the Ca-Cu process in ammonia production plants Isabel Martínez1*, David Armaroli1, Matteo Gazzani2, Matteo C. Romano1 1
Politecnico di Milano, Department of Energy, Via Lambruschini 4, 20156 Milano (Italy) 2
ETHZ, Institute of Process Engineering, Sonneggstrasse 3, 8092 Zurich, Switzerland
*corresponding author:
[email protected] Abstract In this work, the application of a Ca-Cu process into a state-of-the-art ammonia production plant is assessed through process simulations, aiming at reducing the primary energy needs of the syngas production and purification section. The proposed process has shown to significantly reduce the specific primary energy consumption, even when accounting for the higher electric consumption associated to the Ca-Cu process. From an environmental point of view, the proposed process has an inherently high CO2 capture efficiency (about 97%). The encouraging thermodynamic performance along with some simplifications in the syngas conditioning process and the ammonia synthesis loop, confirm the potential of the Ca-Cu process as route for syngas generation in ammonia plants.
Keywords: ammonia production, urea production, sorption enhanced reforming, chemical looping, CCS
Nomenclature ATR
Autothermal reformer
CAPEX
Capital expenditure
CCR
Carbon capture ratio
CCS
Carbon capture and sequestration
CLC
Chemical looping combustion
ECO
Economizer
EVA
Evaporator
FTR
Fired tubular reformer
HP
High pressure
HT
High temperature
HX
Heat exchanger
LHV
Lower heating value
LP
Low pressure
LT
Low temperature
MDEA
Methyldiethanolamine
NG
Natural gas
OPEX
Operating expenditure 1 ACS Paragon Plus Environment
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PSA
Pressure swing adsorption
SER
Sorption-enhanced reforming
SH
Superheater
SMR
Steam methane reforming
SPECCA
Specific primary energy consumption for CO2 avoided
WGS
Water gas shift
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1. Introduction
2
The chemical and petrochemical sector is the largest industrial consumer of energy, accounting for 30% of
3
the total industrial energy use1. Within this sector, ammonia production accounts for 17% of the energy
4
consumed, being the second largest inorganic chemical produced worldwide (137 million tons in 2012)2.
5
Even though ammonia is used in a diversified set of industrial sectors (e.g. as precursor to nitrogen
6
compounds for fertilizers, explosives, fibers and plastics, as refrigerant or as remediation for gas emissions),
7
its principal use nowadays is in the fertilizer industry, with more than 80% of the ammonia produced
8
worldwide dedicated to this sector3. Accordingly, the increase in ammonia production in the last decades has
9
been driven by the demographic growth of developing countries such as China, and it is therefore expected
10
to keep growing in the coming decades. The major ammonia producer worldwide is China, which contributes
11
with more than 32% of the global production, followed by India, US and Russia. While in China the largest
12
fraction of ammonia is produced starting from coal gasification, about 80% of the world production is
13
provided by steam reforming of natural gas. In this case, natural gas is used both as source of hydrogen
14
needed to form NH3, and as fuel for supplying the process energy needs. In such plant, approximately 2/3 of
15
the natural gas introduced is used as a feedstock whereas the remaining input is used for providing the
16
energy needs. As a consequence, ammonia production costs are greatly influenced by the natural gas cost,
17
which may represent up to 70-85% of the final production cost2.
18
Ammonia production consists of three main steps. Firstly, a syngas containing the required stoichiometric
19
ratio between hydrogen and nitrogen for ammonia synthesis (reaction (1)) is produced by steam reforming of
20
natural gas. Although different reforming configurations exist for syngas generation in ammonia plants, the
21
most common layout makes use of a two-step reforming process4. A primary externally-fired tubular
22
reformer is used to convert natural gas in presence of a Ni-based catalyst at around 750-850°C. Energy
23
needed in this primary reforming is supplied by the oxidation of natural gas and wastes from the syngas
24
cleaning. Then, a secondary reforming step, which typically consists of an autothermal reformer (ATR) using
25
air as oxidant, is adopted to introduce the right amount of nitrogen while converting the remaining natural
26
gas load. This secondary reforming step takes place at a higher temperature (around 1000°C) to achieve high
27
conversion of methane. In the second step of the ammonia production process, syngas is conditioned to meet
28
the specifications of the ammonia conversion step: CO is shifted to H2 in a two-step water gas shift (WGS)
29
section, CO2 is removed and, finally, a methanation step reduces the CO content to ppm level. It is however
30
worth mentioning that all main licensors and engineering constructors have developed proprietary solutions
31
for syngas production and conditioning, which also comprise alternative processes as pressure swing
32
adsorption or cryogenic separation4. Finally, in the third main step the purified synthesis gas is compressed
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and sent to the ammonia synthesis reactor, which typically operates at a pressure of around 150-250 bar4.
34 35
N + 3H ↔ 2NH ∆H = −46 kJ/mol
36
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Following the introduction of the disruptive single-train plants by Kellogg in 1963, different modifications
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and advancements to the well-established ammonia production process have been proposed and studied,
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mainly with the aim of (i) reducing the specific energy consumption per unit of ammonia produced and of
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(ii) increasing the plant capacity. Many of the modifications proposed concern the reforming step, which is
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indeed responsible for a large share of the CAPEX and OPEX. Concerning the latter, it is of paramount
42
interest to reduce the amount of fuel needed for sustaining the reforming reactions in the fired tubular
43
reformer, which is conducive to reduction in the primary energy consumption. However, despite the
44
numerous improvements achieved in the last decades, the specific energy requirement for ammonia
45
production is still significantly larger than the theoretical minimum4.
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Despite being the preferred process for syngas production at medium-large scale, steam methane reforming
47
(SMR) of natural gas features important drawbacks: the need to keep the reformer size limited requires to
48
operate the system at moderately high pressure (30-50 bar) while high methane conversion is achieved
49
operating the system at high temperature (methane conversion, which follows an endothermic reaction,
50
approaches equilibrium at the outlet of the reformer). With the aim of overcoming the thermodynamic
51
restrictions imposed by this configuration, multiple studies in the scientific literature present enhanced
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processes, especially where the natural gas reforming is combined with the CO2 separation. The sorption
53
enhanced reforming (SER) concept has received great attention in the past decades due to some inherent
54
advantages over conventional reforming processes5. In particular, the SER process makes use of a CO2
55
acceptor material (usually a Ca-based sorbent) in the reformer, which allows reforming (reaction (2)), water
56
gas shift (reaction (3)) and carbonation (reaction (4)) reactions to occur simultaneously in the same reactor6–
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8
58
towards the hydrogen production, leading to syngas with a hydrogen content higher than 96%vol. (in dry
59
basis) in a single step and resulting in a much intensified process than SMR. Moreover, the energy released
60
by CO2 carbonation drives the reforming reaction, enabling the use of adiabatic reactors, avoiding the
61
necessity of heat transfer at high temperature through pressurized tube walls.
. The CO2 removal from the gas phase pushes the equilibrium of the reforming and water gas shift reactions
62 63
CH + H O ↔ CO + 3H ∆H = +206.2 kJ/mol
(2)
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CO + H O ↔ CO + H ∆H = −41.5 kJ/mol
(3)
65
CaO + CO ↔ CaCO ∆H = −178.8 kJ/mol
(4)
66 67
The main challenge of the SER concept is the regeneration of the CaCO3 formed by the carbonation reaction,
68
which is highly endothermic and requires additional fuel to provide the energy needed. Among the different
69
solutions proposed in literature for supplying the CaCO3 regeneration energy, the configuration that couples
70
an exothermic reduction reaction of a CuO-based material with the calcination step is particularly promising,
71
and it has thus received great attention in the last years9–11. The sequence of reactions in this novel process is
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depicted in Figure 1. The proposed process is carried out in a series of fixed bed reactors, where temperature,
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pressure, feeds and products vary during the different stages of the process. The first step (indicated as A in 4 ACS Paragon Plus Environment
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Figure 1) comprises the hydrogen production through a SER process according to the principles discussed
75
above. The solid bed contains the Ca-based sorbent, the Ni-based reforming catalyst and the Cu-based
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material, all in their reduced form as a result of the preceding reduction/calcination stage of the Ca-Cu
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process. When feeding natural gas and steam, the Ca-based sorbent reacts to form CaCO3, while the
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reforming catalyst and the Cu-based material remain inert. This stage of the process is carried out at high
79
pressure (10-30 bar). During the following step of the process (B in Figure 1), Cu present in the solid bed is
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oxidized to CuO by feeding pressurized diluted air into the reactor. Diluted air enables a fine tuning of the
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temperature and oxygen concentration at reactor inlet, which need to be limited for controlling the maximum
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temperature achieved and so minimize the CaCO3 decomposition. In this step, virtually all the O2 in the inlet
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air reacts with the oxygen carrier and the resulting off-gas is high purity N2 with small amounts of CO2 and
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Ar. In the last step (stage C), the calcination of the CaCO3 formed during stage A is driven by the exothermic
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reduction of the CuO, which is carried out feeding a mixture of CH4, CO and H2 into the reactor. With the
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aim of limiting the calcination temperature needed, this last stage of the process operates at around
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atmospheric pressure. The Ca/Cu molar ratio in the solid bed is chosen for ensuring that during stage C the
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heat released in the CuO reduction is enough for sustaining the endothermic CaCO3 calcination.
89 90
[FIGURE 1]
91 92
A full process design of a stand-alone hydrogen production plant based on this process was investigated by
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Martinez et al.11, who showed that a Ca-Cu process can produce H2 with high efficiency while also capturing
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more than 90% of the produced CO2. However, the Ca-Cu process has never been investigated as part of the
95
syngas generation and purification in an ammonia production plant, despite the inherent advantage of
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providing almost pure streams of hydrogen and nitrogen as part of its products. Accordingly, this work
97
addresses the integration of the novel Ca-Cu looping process within an ammonia production plant. A case
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study of a full system integration of the Ca-Cu process into an ammonia production plant is presented.
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Modifications of the ammonia synthesis loop when integrated with this concept are investigated with the aim
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of assessing the impact from a full process performance point of view. The performance obtained for the
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novel Ca-Cu based ammonia plant is compared with a benchmark ammonia plant based on the state-of-the-
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art technologies.
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2. Reference ammonia production plant
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A state-of-the-art, 1500 tons/day, ammonia plant is considered as benchmark for comparison with the
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advanced ammonia production process integrated with the Ca-Cu concept analyzed in this work; to this end,
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mass and energy balances of the syngas production and ammonia synthesis area in the reference plant were
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solved. A schematic of the reference ammonia production plant, which represents the most widespread
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configuration for ammonia production, is depicted in Figure 2 and Figure 3. The following three processing
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sections characterize the reference ammonia production plant: (i) the synthesis gas production island, 5 ACS Paragon Plus Environment
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covering natural gas reforming processes as well as synthesis gas conditioning and purification steps (i.e.
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water gas shift reactors, CO2 removal and methanation stages), (ii) the ammonia production loop, and (iii)
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steam cycle (not shown in the figures) fed by steam produced by recovering the heat available in the plant.
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These balances have been solved using Aspen Plus. For the synthesis gas production island, the cubic
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equation of state of Redlich-Kwong-Soave (RKS)12 has been used, whereas for the ammonia synthesis loop
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the RKS equation modified by Mathias et al. (1984)13 has been preferred due to its suitability in predicting
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real gas behavior. For pure water and steam, thermodynamic properties from the International Association
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for the Properties of Water and Steam (IAPWS)14 have been used. Tables S1 and S2 of the Supporting
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Information summarize the main assumptions used for calculating the synthesis gas production island and
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the ammonia production loop, respectively, in the reference ammonia plant.
121 122
[FIGURE 2]
123
[FIGURE 3]
124 125
Catalyst used for steam reforming processes is highly sensitive to any sulphur compound. Therefore, the
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concentration of these compounds has to be below 0.15 ppm before feeding the NG into the reforming
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reactor. Typically, a non-regenerative ZnO-based solid bed is used for removing sulphur compounds in a
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two-stage process. Sulphur compounds are first converted into hydrogen sulphide (H2S) by hydrogenation in
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a catalytic solid bed made of Co and Mo oxides supported on alumina at temperatures not higher than 340-
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370°C15. A H2 concentration of around 2%vol. is needed at the inlet of this hydrogenation step for ensuring
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that H2S formation proceeds at sufficient rate16; thus, part of the hydrogen leaving the PSA unit adopted in
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the synthesis section of the plant is recycled to the feed. Once sulphur compounds are in the form of H2S, the
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gas is fed to the ZnO-based solid bed that operates at the same temperature of the hydrogenation process. In
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this reactor, H2S reacts with ZnO to form ZnS, and the H2S concentration in the gas phase drops below 0.1
135
ppm15.
136
After desulphurization, NG is mixed with high pressure steam produced within the plant and fed into an
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adiabatic pre-reformer. The steam-to-carbon (S/C) ratio typically employed in ammonia production plants
138
ranges between 2.5 to 3.6, with a growing trend towards the lower bound in the most modern plants17. In this
139
work, an S/C ratio of 3 has been chosen for the reference plant. The pre-reforming stage enables the
140
decomposition of higher hydrocarbons in the NG into CO and H2 before entering the primary reforming
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section, where the high temperature may cause solid carbon formation from these hydrocarbons and
142
consequent catalyst deactivation. In addition, the use of this pre-reformer reactor reduces the reforming duty
143
needed in the subsequent primary reforming section. Pre-reforming is typically carried out in an adiabatic
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reactor with an inlet gas temperature of around 380-500°C in the presence of a Ni-based catalyst18. For the
145
reference plant, a pre-reformer inlet temperature of 490°C has been chosen. Pre-reformed gas is then heated
146
up to 620°C before being fed to the primary reforming reactor.
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The primary fired tubular reformer (FTR) consists of tubes filled with catalyst that are placed inside a
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furnace where CO/H2/CH4-rich off-gas from a downstream PSA unit and additional NG are burnt with air for
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supplying the energy needed for the reforming. Typically, CH4 conversions in the range of 55-65% are
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achieved in the primary reforming process working with outlet temperatures of 780-830°C and leading to
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CH4 concentrations in the product gas of 10-15% on dry basis4,17. Such values of CH4 conversion at FTR
152
outlet are below the conversion values corresponding to equilibrium at those FTR outlet temperatures. In the
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reference plant considered in this work, FTR outlet conditions of temperature of 820°C and methane
154
conversion 58%, have been assumed, which lead to a CH4 content in the product gas of 12.5% on dry basis.
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The sensible heat of the flue gas, which exits the FTR furnace at 1010°C, is recovered heating up the FTR
156
charge fed to the primary reforming section and the air streams needed in the FTR furnace burners and in the
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secondary reformer. Secondary reforming process is based on an autothermal reformer where air is fed
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together with the partially reformed gas from the FTR. The assumed ATR outlet temperature of 1000°C
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results into a residual CH4 content lower than 0.5% on dry basis, calculated based on equilibrium. Nitrogen
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introduced with air in the ATR provides the N2 needed for the ammonia synthesis reaction (reaction (1)),
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which should be fed with a gas containing a H2/N2 ratio of 3. Methane conversion value of 58% in the FTR is
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calculated so that the amount of air needed to close the energy balance in the ATR allows fulfilling the H2/N2
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molar ratio needed at the ammonia synthesis reactor inlet. Another parameter to be defined is the air feeding
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temperature into the ATR, which is usually between 520°C and 610°C; in the reference plant in this work,
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the air is at 540°C. A three-stage intercooled compressor (with isentropic and mechanical-electric
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efficiencies of 75% and 98%, respectively) is used to increase the air pressure to about 33 bar at the ATR
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inlet.
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Syngas obtained from ATR contains almost 15% of CO (in dry basis), which is converted into H2 and CO2
169
via the WGS reaction (3). The WGS section consists of two stages with intermediate cooling. High
170
temperature WGS operates with a syngas inlet temperature of 350°C using an iron/chromium oxide based
171
catalyst, whereas the second WGS step operates at a lower temperature of 200°C using a copper/zinc oxide
172
based catalyst. At the outlet of the WGS section, CO content in the syngas is reduced to less than 0.4% (dry
173
basis), whereas CO2 content is around 18% (dry basis). The syngas from the WGS section is cooled down to
174
35°C, which allows removing most of the water by condensation, and is sent to the CO2 removal section.
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Multiple CO2 removal solutions exist, depending on the vendor and the specifications of the downstream
176
ammonia process: here, the adoption of a state-of-the-art MDEA scrubbing process is considered. High CO2
177
separation efficiencies can be achieved with this absorption process, while recovering a CO2 stream at low
178
pressure with a purity higher than 99%. Once compressed, this stream is therefore suitable for use in other
179
processes (e.g. urea production) or for sequestration, in case of carbon capture and storage configuration. A
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CO2 separation efficiency of 98.5% is assumed for this MDEA process, with a regeneration energy duty of
181
0.955 MJ/kg of CO219.
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Following the CO2 removal, CO2 and CO content in the syngas is about 0.6% vol., which, along with H2O,
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used in the ammonia synthesis deactivates in the presence of oxygen compounds). Therefore, a methanation
185
step is adopted to convert the remaining CO and CO2 into CH4 and H2O. To this end, syngas exiting the
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MDEA unit is heated up to 280°C and fed to an adiabatic methanation reactor, where a nickel-based catalyst
187
is used to promote the reverse of the reforming and WGS reactions (i.e. reverse of reactions (2) and (3)). H2O
188
is ultimately removed by condensation. Despite the fact that some H2 is consumed through this methanation
189
process, the low oxygen content achieved with the methanation (below 10 ppm) makes this process the
190
preferred one. Due to the exothermicity of the methanation reaction, gas temperature typically increases by
191
50-70°C. The outlet gas is then cooled down to 30-35°C to condensate the H2O, compressed and sent to the
192
ammonia synthesis section.
193
Typically, in modern ammonia plants, centrifugal compressors are used for pressurizing the syngas to the
194
required operating conditions at ammonia synthesis reactor inlet (i.e. 150-250 bar, 350-550°C)20. In the
195
reference plant, a three-stage intercooled compressor is used with intermediate condensate removal steps. As
196
highlighted in the compression train shown in Figure 3, syngas is mixed first with hydrogen (stream #30),
197
which is recovered from the ammonia synthesis waste via a dedicated PSA, and secondly with the recycle
198
gas of the ammonia synthesis loop (#29). Both recycles make the process more efficient, improving the
199
productivity of the plant, while also further reducing the concentration of oxidized species in the feed of the
200
ammonia reactor by dilution.
201
Syngas conditions at ammonia reactor inlet (#19) are 320°C and 199 bar. The outlet gas composition is
202
governed by the reaction rate of ammonia formation (reaction (1)), which at the industrial conditions of
203
interest (i.e. 450°C, 150-250 bar and 10% of inerts at inlet) is limited by the unfavorable equilibrium and by
204
the effect of the operating parameters. Based on the operating conditions of modern ammonia plants, the
205
ammonia content in the product gas (#20) is 20% for the assumed operating conditions4. This value of
206
ammonia content at reactor outlet corresponds to an ammonia conversion of 34% per pass, which
207
corresponds to the conversion at chemical equilibrium calculated at 465°C (15°C above the reactor outlet
208
temperature). For calculating the ammonia synthesis reactor in the reference plant, the aforementioned
209
temperature approach has been chosen for the given outlet temperature of 450°C.
210
Due to the rather low ammonia conversion per pass, unreacted gas at synthesis reactor outlet is sent back to
211
the reactor after condensation of the produced ammonia17. As shown in Figure 2, a purge stream is needed to
212
avoid that inerts (CH4 and Ar) accumulate in the ammonia synthesis loop. Typically, a small fraction of
213
recycled gas (1-2% of the total recycle) is purged (stream #33), which is enough for maintaining a constant
214
fraction of inerts in the synthesis loop of around 10-15%20. This purge gas is first scrubbed with water to
215
remove ammonia, and then sent to a PSA unit for recovering H2, which is sent back to the compression train
216
of the synthesis reactor (#30). PSA off-gas, which consists of H2, CH4, Ar and N2 (#35), is used as fuel in the
217
burners of the primary reforming section. In order to limit the ammonia content in the synthesis gas, which
218
hinders the reaction equilibrium, most of the ammonia present in the converter product is separated by
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condensation at about -20°C (#27). In modern plants, ammonia-rich gas is first cooled down with the
220
recycled gas and then with cooling water to a temperature of around 30-45°C. Afterwards, a refrigeration 8 ACS Paragon Plus Environment
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cycle using ammonia as refrigerant fluid is used for reducing the ammonia-rich gas temperature down to -
222
20°C. This refrigeration cycle operates between 1.5 and 18 bar, which correspond to an evaporation and
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condensation temperatures of the refrigerant fluid of -25°C and 38°C, respectively. Liquid ammonia
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separated is flashed to around 20 bar in a let-down vessel to release dissolved gases (#28) that are sent to the
225
scrubbing and PSA units.
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As anticipated before, a large amount of energy is available by cooling (i) the flue gas of the primary
227
reformer, and (ii) the syngas after the secondary reformer, the WGS and the methanation reactor. Energy is
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recovered by producing high pressure steam at 510°C and 110 bar. This steam is expanded in a turbine that
229
produces electric power for (i) the syngas compressor, (ii) the refrigeration cycle, and (iii) the air compressor
230
for the ATR. The steam required to meet the S/C of the reforming section is provided by bleeding steam
231
from the turbine at the corresponding pressure (33 bar). As shown in the results section, there is an excess of
232
electricity in the reference ammonia plant, which is considered to be exported to the electric grid. Table 1
233
shows the thermodynamic properties and the mass balance of the reference ammonia production plant as
234
numbered in Figure 2 and Figure 3.
235 236 237
Table 1 Temperature, pressure, flow rates and composition of the main process streams shown in Figure 2 and Figure 3 for the reference ammonia production plant
No
T [°C]
P [bar]
G [kg/s]
1
15.0
33.1
2
365.0
3
.
N
Molar composition [%]
[kmol/s]
Ar
CH4
CO
CO2
H2 O
H2
C+
N2
O2
NH3
10.5
0.58
--
89.0
--
2.0
--
--
8.1
0.9
--
--
32.7
8.2
0.46
--
87.2
--
2.0
--
2.0
8.0
0.9
--
--
490.0
32.4
34.9
1.95
--
20.8
--
0.5
76.1
0.5
1.9
0.2
--
--
4
438.5
32.3
34.9
2.03
--
21.9
0.3
2.4
69.2
6.2
--
0.2
--
--
5
620.0
32.1
34.9
2.03
--
21.9
--
2.4
69.2
6.2
--
0.2
--
--
6
352.0
1.2
48.4
1.67
1.0
--
--
--
--
--
--
78.0
21.0
--
7
538.0
32.7
19.4
0.67
1.0
--
--
--
--
--
--
78.0
21.0
--
8
1010.0
1.2
51.6
1.84
1.2
--
--
8.3
15.9
--
--
71.7
3.0
--
9
820.0
31.8
34.9
2.54
--
7.4
5.8
6.2
40.7
39.7
--
0.2
--
--
10
1000.0
31.3
54.3
3.43
0.2
0.2
9.3
4.9
32.7
37.4
--
15.3
--
--
11
350.0
31.3
54.3
3.43
0.2
0.2
9.3
4.9
32.7
37.4
--
15.3
--
--
12
428.1
31.1
54.3
3.43
0.2
0.2
2.2
12.0
25.7
44.4
--
15.3
--
--
13
200.0
30.7
54.3
3.43
0.2
0.2
2.2
12.0
25.7
44.4
--
15.3
--
--
14
223.8
30.5
54.3
3.43
0.2
0.2
0.2
14.0
23.6
46.5
--
15.3
--
--
15
35.0
30.1
39.7
2.62
0.3
0.2
0.3
18.3
0.2
60.7
--
20.0
--
--
16
280.0
29.9
18.9
2.15
0.3
0.3
0.3
0.3
0.2
74.1
--
24.4
--
--
17
35.0
29.4
18.5
2.10
0.3
0.9
--
--
0.2
73.6
--
25.0
--
--
18
144.7
200.0
65.4
6.66
1.8
7.2
--
--
--
66.8
--
22.6
--
1.6
19
320.0
199.0
65.4
6.66
1.8
7.2
--
--
--
66.8
--
22.6
--
1.6
20
450.0
196.0
65.4
5.64
2.1
8.5
--
--
--
51.8
--
17.6
--
20.0
9 ACS Paragon Plus Environment
Industrial & Engineering Chemistry Research
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21
76.7
193.0
65.4
5.64
2.1
8.6
--
--
--
51.8
--
17.6
--
20.0
22
30.0
192.0
55.7
5.08
2.3
9.3
--
--
--
57.4
--
19.5
--
11.4
23
30.0
192.0
9.6
0.57
0.5
1.4
--
--
--
1.1
--
0.4
--
96.6
24
-20.0
190.0
55.7
5.08
2.3
9.3
--
--
--
57.4
--
19.5
--
11.4
25
-20.0
190.0
47.5
4.59
2.5
10.2
--
--
--
63.4
--
21.6
--
2.3
26
25.0
189.0
47.5
4.59
2.5
10.2
--
--
--
63.4
--
21.6
--
2.3
27
-20.0
190.0
8.2
0.48
0.5
1.1
--
--
--
0.4
--
0.1
--
97.9
28
40.0
190.4
0.5
0.03
11.4
32.0
--
--
--
27.4
--
9.2
--
19.9
29
146.7
189.0
46.9
4.53
2.5
10.2
--
--
--
63.4
--
21.6
--
2.3
30
20.6
58.0
0.1
0.03
--
--
--
--
--
100
--
--
--
--
31
25.3
20.0
17.4
1.02
0.2
0.4
--
--
--
--
--
--
--
99.4
32
25.0
189.0
0.6
0.06
2.5
10.2
--
--
--
63.4
--
21.6
--
2.3
33
25.4
189.0
1.1
0.09
5.4
17.5
--
--
--
51.5
--
17.5
--
8.2
34
20.8
60.0
1.0
0.08
5.9
19.0
--
--
--
56.0
--
19.0
--
0.1
35
20.7
58.0
0.9
0.04
11.9
38.3
--
--
--
11.3
--
38.4
--
0.2
238 239
3. Ammonia production plant integrated with the Ca-Cu process
240
In this section, the different assumptions and process conditions chosen for solving mass and energy
241
balances of the ammonia production plant integrated with the Ca-Cu process are described. This plant
242
includes the following sections: (i) sulphur removal and pre-reforming of natural gas, (ii) production of
243
syngas through the Ca-Cu process, (iii) syngas compression train, (iv) ammonia synthesis reactor, and (v)
244
refrigeration cycle of the ammonia synthesis loop. Table S3 of the Supporting Information summarizes the
245
main process assumptions used for calculating the Ca-Cu process as well as the heat recovery steam cycle,
246
whereas the conditions used for the ammonia synthesis and refrigeration systems correspond to those given
247
in Table S2. Natural gas composition and desulphurization unit temperature used in this case are equal to
248
those shown in Table S1 for the reference plant. As in the reference ammonia plant, balances have been
249
calculated for a NH3 production of 1500 tons/day. Regarding the materials for the Ca-Cu process, CaO-based
250
sorbent composition considered is 85%wt. of CaO over an inert support (considered as Al2O3 for the
251
calculations) with a CaO conversion of 40%, which has been set based on the literature found on high
252
stability CaO sorbents21,22. For the Cu-based material, a high Cu content material is considered (i.e. 65%wt.
253
of CuO over Al2O3) that has been proven to be resistant towards long term agglomeration and deactivation
254
for conditions similar to those used in the Ca-Cu process23.
255 256 257
3.1. Sulphur removal and pre-reforming of natural gas
258
As in the reference ammonia production plant, natural gas needs to be desulfurized to avoid poisoning of the
259
reforming catalysts. Therefore, the same desulfurization process of the reference ammonia production plant
260
is considered. After desulfurization, the natural gas is mixed with high pressure steam, pre-heated, and fed to 10 ACS Paragon Plus Environment
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Industrial & Engineering Chemistry Research
261
the pre-reforming unit. The main objective of this unit is to convert the higher hydrocarbons present in the
262
natural gas into CO and H2, since these hydrocarbons could be degraded into coke via thermal cracking at the
263
high temperatures in the reforming stage of the Ca-Cu process. In case coke formation occurs, it would
264
deposit over the material particles present in the solid bed, which may cause particle breakup, deactivation or
265
hot spots due to combustion in the following stages. Accordingly, two heated pre-reformers are adopted
266
upstream of stages A and C of the Ca-Cu process, working at high and low pressure respectively, where
267
chemical equilibrium is achieved. The heat needed for the reforming reactions in this kind of reactors is
268
supplied by cooling a high temperature flue gas24. From the point of view of the Ca-Cu process operation, the
269
use of the heated pre-reformers in this plant instead of the adiabatic pre-reformer of the reference plant leads
270
to higher conversion of hydrocarbons into CO, CO2 and H2 and the achievement of chemical equilibrium at
271
the Ca-Cu reactors inlet temperature. As a consequence, the possibility of CaO hydration is reduced and the
272
temperature profiles in the bed are stabilized. As a matter of fact, solid bed temperature in the initial section
273
of the packed bed reactor under stages A and C will remain stable at the temperature of the gas from the
274
corresponding heated pre-reformer placed upstream. The main operating conditions of the pre-reformers
275
have been indicated in Table S3 of the Supporting Information. In particular, the operating temperature of
276
the high pressure pre-reformer before stage A is 700°C, which has been chosen to ensure that conditions
277
within the reactor during stage A make CaO hydration impossible based on thermodynamics (i.e. that
278
pH2O
