Thursday, 9 February 2017

Energy Conservation in CO2 Removal System of Ammonia Plant

Energy Conservation in CO2 Removal System of Ammonia Plant
IFFCO, Phulpur$FILE/Article.pdf

The selection and design of CO2 system was the most difficult engineering job of the Phulpur
Expansion Project (PEP). Though a repeat of Aonla Expansion Project (AEP), the use of 100
percent naphtha feed in place of 50 percent naphtha + 50 percent natural gas feed ( as in AEP )
would have resulted in higher CO2 and a change in raw gas composition. Consequently, a
thorough study was required to check the adequacy of this section. Further, the new ammonia
plant was consuming higher energy per tonne of ammonia plant was consuming higher energy
per tonne of ammonia as compared to the design value and the CO2 removal system was
identified as one of the higher energy consuming areas. This paper gives an account of the study
carried out to check the adequacy of the CO2 removal system, the changes incorporated in this
system and the efforts made to lower the energy consumption of CO2 removal
Indian Farmers Fertiliser Cooperative Ltd., ( IFFCO ), with an annual production of over 40 lakh
tonnes of fertilisers is considered to be one of the largest manufacturer of fertiliser in the world.
During December 1997, IFFCO installed an expansion project comprises of 1350 MTPD
ammonia plant designed and engineered by M/s. HTAS and a 2200 MTPD urea plant designed
and engineered by M/s. Snamprogetti.
The Phulpur Expansion Project (PEP) is a repeat of Aonla Expansion Project (AEP). However,
due to change in geographical layout, most of the battery limit engineering had to be re-done.
Since PEP was based on the use of 100 percent naphtha feed in place of 50 percent naphtha and
50 percent natural gas feed (AEP), a new per-hydrodesulphurisation section was added at PEP
and certain sections of plants had to be rechecked for adequacy. The sections rechecked for
adequacy are (i) steam generation and boiler feed water and process condens ate systems, since
naphtha plants require higher quantity of steam than natural gas (ii) synthesis gas compressor
turbine for higher extraction of medium pressure steam and CO2 removal system, since higher
quantities of CO2 have to be handled in a naphtha based plant.
Ammonia plant - The Latest Generation Plant
Some of the salient features for energy saving in ammonia plant are as follows:
-Gas turbine drive, with naphtha as fuel, for process air compressor
-Heat recovery unit connected to the gas turbine for generating high pressure steams to meet the
requirement of ammonia and urea plants.
-Medium pressure process condensate stripper
-GV - Low energy CO2 removal system
CO2 Removal Systems - Problems and Solution
The selection and design of CO2 removal system was the most difficult engineering job of this
project. Though it was a repeat of AEP, the use of 100 percent naphtha feed in place of 50
percent naphtha and 50 percent natural gas feed (used in AEP), would have resulted in higher
quantity of CO2 generation and a change in raw syn gas composition. Consequently a thorough
study was carried out to check the adequacy of this system like the adequacy of the existing
towers, the performance of the repeat CO2 blower, the stripper top overhead condenser meeting
the duty with respect to CO2 temperature and cooling water flow rate and the GV pumps giving
the required solution flows etc.
CO2 absorption is more in case of naphtha feed as compared to gas feed, because of higher
partial pressure of CO2 in feed gas. Though the solution flow rates are higher the CO2 absorbed
(Vol. flow) per unit volume of solution is higher in case of naphtha feed and hence the same size
of absorber could be used as that of AEP.
The main change envisaged was the change of packing in the absorber bottom bed lower part
from IMTP-50 to CMR - 3 (Cascade Mini Rings) due to increase in pressure drop because of
higher gas flow rates. These packings have same
efficiency as IMTP but offer lower pressure drop. Also the lower layer of the 3rd bed packing
have been changed from IMTP - 40 to IMTP-50. The packing volumes remained unchanged with
same diameter of absorber. The distributors also remained the same.
For regeneration, the CO2 liberation due to flashing is more and so the capacity (total tower size)
of the regenerators were checked and found suitable.
The process control principles as well as solution composition and physico - chemical
characteristics were also checked and found suitable.
CO2 blower capacity was checked with the vendor. The capacity of the CO2 blower was
increased as it had to handle a higher CO2 flowrate.
Stripper top CO2 overhead condensers capacity with respect to CO2 cooling load, cooling water
supply, pipe size etc. were also checked and found to be suitable.
The lean, semilean pumps and the hydraulic turbines were also checked and found to be adequate
without any modification.
CO2 Removal System and its Chemistry
In the GV CO2 removal system, carbon-di-oxide is removed from the process gas by absorption
in an aqueous hot potassium carbonate solution (HPC solution) containing approx. 30 wt. percent
potash (K2CO3) partly converted into bicarbonate (KHCO3). The solution also contains glycine
and a tertiary amine DEA as activator and Vanadium Penta Oxide as corrosion inhibitor. The
reason for keeping the solution hot is to increase the rate of absorption and keep the bicarbonate
A scheme of CO2 removal system is illustrated in Figure 1.
The process gas from the shift reactors is passed to the CO2 absorber (F 3303). In the absorber
the gas flows upwards against a descending stream of hot potassium carbonate solution.
Approximately 15 percent of the solution is introduced above the top bed at 70 deg. C whereas
the remainder is introduced at about 106 deg C below the two top beds.
The CO2 absorption occurs according to the following reaction mechanism:
CO2+H2O = HCO3+H .... (1)
CO3+H2O = HCO3+OH ..... (2)
CO3+CO2+H2O = 2HCO3 .... (3)
The reaction rate of step 3 is determined by step 1 which is the slower of the steps 1 and 2. The
activator action resulting in an increased rate is caused by the quick transfer of gaseous CO2 into
the liquid phase by means of the glycine carbamate formation according to the reaction.
The activator effect is much higher than the one relating to carbamate concentration . At higher
temperature and in the presence of OH the carbamate
is hydrolysed and the activator is restored according to the reaction :
The sum of steps 4 and 5 gives 1. As reaction 4 and 5 takes place continuously, it means that the
glycine acts as a CO2 carrier. Reaction 5 is the hydrolysis of glycine carbamate. The reaction is
catalysed by a small amount of DEA in the solution .
The absorption takes place in two stages in F 3303. In the first stage, where the bulk of CO2 is
absorbed, the high temperature increases the reaction rates of 5 and 3. This is done by using the
normal regenerated solution (semi lean) from the second regenerator, F3302.
In the second stage, a stream of strongly regenerated solution (lean solution) is utilised. At the
lower temperature, the CO2 vapour pressure of the solution is further reduced to meet the low
CO2 slippage in the purified gas (about 0.05 mole % dry CO2).
The solution leaving the absorber bottom is loaded with CO2 and is referred to as the rich
solution. The rich solution is depressurised through the hydraulic turbines, TX 3301 A/B driving
the semilean solution pumps P 3301 A/B.
Downstream the hydraulic turbines, the rich solution enters the top of the first regenerator, F3301
working at a pressure of about 1.0 Kg/cm2g.
A stream of rich solution extracted from the top of F3301 is depressurised through a control
valve and enters the top of the second regenerator, F 3302, working at a low pressure (0.1
Kg/cm2g). The semilean solution extracted from the intermediary tray of F3301 at 124 deg C is
flashed across the level control valves and fed into F 3302. In F 3302 the semilean solution is
collected on a take-off tray feeding the semilean pumps, P 3301 A/B/C. The lean solution
extracted from bottom of F3301 at 127 deg. C is flashed across the level control valves and
enters F 3302 below the semilean solution take-off tray feeding the semilean pumps, P 3301
A/B/C. The lean solution extracted from bottom of F3301 at 127 deg. C is flashed across the
level control valves and enters F 3302 below the semilean solution take off try. The lean solution
is collected in the bottom of F3302 and fed to the lean solution pumps, P3302 A/B. The steam
developed by flashing of the lean and semilean solutions is used as stripping steam to regenerate
the rich solution entering the top of F 3302. In this way more than half of the stripping steam
supplied to F3301 is recovered by flashing and works in double effect in F 3302. The semilean
solution is pumped by P3301 A/B/C and sent to the lower part of the absorber (F3303) at a
temperature of 106 deg. C Lean solution is drawn from the bottom of F 3302 at a temperature of
109 deg C and cooled by the DMW preheater E 3306 A/B to a temperature of 70 deg C. From E
3306 A/B, the cooled lean solution is fed into the top of F 3303 by P 3302 A/B.
The gas cooling arrangement is illustrated in Figure 2.
Before absorption of carbon dioxide, the shifted process gas is cooled from about 163 deg C to
about 104 deg C.
First step is cooling from about 163 deg C to 132 deg C in the parallel operated vetrocoke
reboilers, E 3302 A/B, and thereby providing a part of the heat required for the first regenerator,
F 3301.
The process gas from E 3302 A/B is then cooled from about 132 deg C to 122 deg C in the LP
boiler, E 3303. The steam produced in E 3303 is by means of the ejector, X 3301, injected into
the first regenerator, F 3301 as stripping steam.
The process gas is further cooled to 104 deg C in the DMW preheater, E 3304 before the process
gas enters the CO2 absorber, F 3303.
Mechanical Filtration
The purpose of mechanical filtration is to keep the solution clean and free from solids. These
solids may be impurities of the chemicals used for solution preparation and make-up, catalyst
dust entrained by the process gas, rust from steel equipment particularly after a long shutdown
period or fine dust from poor quality activated carbon. Impurities may also enter the sump tank
Two mechanical filters are provided in the plant. The main filter is placed upstream of the
activated carbon filter in order to protect carbon filter from solid build-up. The second filter (of
smaller surface) is placed downstream of the activated carbon filter in order to retain any fine
dust released from poor quality or crashed carbon which can seriously affect the plant's smooth
operation. The filtering medium is polypropylene (alkaline resistant) for temperature not
exceeding 100 deg C. the solution fed to the main filter is about 5% of the circulating flow while
only 50 percent of the solution exiting the filter feeds the activated carbon filter and the
downstream mechanical filter.
Activated Carbon Filtration
This filtration is intended for removing any organic substances (active as well as degraded
organic activators) from the solution. These substances such as grease, lubricating oil, paints,
bitumen, epoxy resins degraded organic compound etc., are saponifiable by the alkaline solution.
They may cause foaming problems and consequent worsening of the CO2 removal plant's
performance. The carbon filter treats the lean and semilean solution streams in similar way. It is
placed downstream of the mechanical filter for extending the carbon life. Only a fraction (about
50 percent) of the solution coming out of the mechanical filter feeds the activated carbon filter.
The most advisable temperature range for the filtration is 70-110 deg C.
Energy Reduction in GV Section
Whatever calories are used in the fertiliser plant, the lion share is always consumed by the
ammonia plant. The new ammonia plant has been running at an energy level higher than the
design value. A task force was formed to identify the areas of high energy consumption in
ammonia plant and it was found that CO2 removal system was one of the areas where the energy
consumption was very high.
The CO2 removal system designed by M/s. Giammarco Vetrocoke, Italy, has the following
guaranteed parameters at the battery limit.
a) 0.05 mole% CO2 (max.) on dry basis at the gas absorber outlet.
b) The CO2 product of 45792 Nm3/h (minimum) at 100% ammonia plant load on dry basis
coming from solution regeneration will be available at the battery limit.
c) The pressure of above said CO2 will be 0.6 Kg/cm2 g minimum.
d) The purity of the above said CO2 will be minimum 99% (vol.) on dry basis.
e) The LP steam consumption shall not be more than 9000 Kg/h with a total specific regeneration
heat consumption of 660 Kcal/Nm3 of CO2.
During the ammonia plant guarantee test run, all the parameters were within the guaranteed
values except for higher LP steam consumption and hence higher regeneration heat.
Complete operating data, solution analysis, gas analysis etc. were carefully studied and analysed.
Following observations were made.
1. The pressure of LP boiler E 3303 was not coming down to the design value of 0.86 Kg/cm2 g
and always above 1.0 Kg/cm2g even with innumerable manipulations of ejector X 3301
operation even with innumerable manipulations of ejector X 3301 operation.
2. The CO2 slip from bottom two beds of absorber F3303 was very much within the design
limits, but the CO2 slips from top two beds were not upto the mark, though the final CO2 slip
from absorber was less than 500 ppm.
3. The LP stripper F 3302 was suspected to be performing below its design performance as
indicated by higher fractional conversion in lean solution than design.
These observations indicated that ejector was not performing and the suction steam quantity was
less and therefore more LP steam was being put in the stripper to meet the solution regeneration
heat requirement. The absorption qualities of semilean solution or performance of bottom beds of
absorber was excellent, while the absorption qualities of lean solution or absorber top beds
performance was bad. The LP stripper was not performing as per design.
The regeneration heat in CO2 removal system was very high. LP steam to the HP regenerator of
GV system was to the tune of 24,000 Kg/h. Any attempt to reduce the steam was causing high
CO2 slip and consequent increase in methanator temperature and loss of valuable potential
The process parameter were tuned to the optimum level like :
1. The solution flow rates were mimimised.
2. The absorber gas inlet temperature was reduced.
3. The semilean solution flow from HP stripper to LP stripper was brought to design value.
With all these efforts the LP steam consumption could be brought down to the tune of 20.0 t/h.
After this, it was decided to check the concentration of chemicals in GV solution.
Through investigation revealed that actual active quantity of the activators (DEA and Glycine)
was less than the design values. Since no standard analytical method is available for quantifying
the degraded quantity of glycine and DEA, and also the standard analysis can give only the total
DEA and glycine (active as well as degraded) there was a chance that the actual active quantities
of activators may be running low. It was therefore essential to remove the degraded activators.
The degraded organic product could be removed only by taking the activated carbon filter in line.
But there was problem in taking the carbon filter in line as the mechanical filter downstream of
carbon filter was not working and hence the problem was getting aggravated.
The activated carbon filter and its downstream mechanical filter were taken in line and partial
solution was filtered for 24 hrs. to remove the degraded organic compounds of activators. After
filtration the foaming of solution was reduced considerably.
In the next step the solution composition was checked.
The GV operating manual reports the average composition of the circulating solution. The
composition does not refer to any specific zone in the plant. Yet, on the grounds of the lab
analysis at lean and semilean pump suction, the average value can be assessed.
As a rule, the concentration of K2CO3, KHCO3, V2O5, V5 Glycine, DEA, Chloride and Fe is
analysed in the sample withdrawn at lean pump suction. While only K2CO3 and KHCO3 is
determined in the sample withdrawn at semilean pump suction.
The fractional conversion and equivalent K2CO3 of each sample can be inferred from the values
of K2CO3 and KHCO3. From the K2CO3 equivalent of each sample it is possible to determine
the K2CO3 equivalent of the overall solution by means of the weighted average of the relevant
flow rates.
Finally, from the K2CO3 equivalent in the overall solution and the K2CO3 equivalent of lean
solution; and the concentration of Glycine, DEA, V2O5 etc. in the lean solution, it is possible to
determine the concentration of th e Glycine, DEA, V2O5 etc. in the overall solution as shown in
the example given below.
As can be seen from the given example that though the analysis of solution indicates a high
percentage of DEA and Glycine (0.9% Glycine and 1.12% DEA), in fact their concentration is
very low i.e. DEA 0.93 % and Glycine 0.75 % only.
After achieving the guaranteed steam consumption, it was decided to have a closer look at the
performance of ejector X 3301 and performance of LP stripper F 3302.
Due to the poor performance of the ejector, the pressure in LP boiler E 3303 was much higher
than design and the process gas exited LP boiler E 3303 at much higher temperature than design.
This means that the process gas heat
transferred to the CO2 removal section was less than design and therefore less heat of
regeneration was being provided to GV solution through process gas. The process gas heat which
should have gone to CO2 removal system was getting transferred to DM water heating and,
therefore, excess external energy (in the form of LP steam) had to be provided in CO2 removal
system to meet the regeneration requirement. Operating data around the ejector under various
conditions were collected and communicated to the ejector designer in (Table 1).
Following facts were established based on the analysis and perusal of operating parameters
around LP stripper F 3302.
1. The Fractional conversion of 0.12 of the lean solution exit HP stripper F 3301 was much better
than the FC of 0.26 of lean solution at the exit of LP stripper F3302.
2. Value LV-9 for regulation of the lean solution between HP and LP strippers was open by 25%
(a very low value which is not found in similar plants), whereas valve LV-7 for regulation of the
semilean solution between HP and LP strippers was fully open.
3. Compared to the first stage (bottom two bed) exit (CO2 slip of <5000 ppm), the CO2 slip at
Absorber 2nd stage (top two bed) exit was too high (450 ppm).
4. There is a difference of only 2 deg. C (vs 3 deg C of design) in the boiling temperature of lean
and semilean solutions from the LP stripper.
All the above observations have pointed towards leakage of semilean solution in the LP stripper
to the lean solution. The the LP stripper to the lean solution. The leakage point appears to be the
semilean solution draw off tray. The stoppage of the leakage will improve the final CO2 slip
from the absorber, which is expected to decrease to about 200 ppm. This will result in
remarkable increase of ammonia production with same process feed. Moreover, it will be easier
to correct the solution flowrates and the valve on the lean solution piping between HP and LP
stripper will sensibly open allowing to control the levels better.
After having performed the above studies, the operating parameters around CO2 removal system
were collected and the regeneration energy was calculated. The calculated regeneration energy
came down from more than 800 Kcal/Nm3 of CO2 to approximately 661 - 674 Kcal/Nm3 of
CO2 remove. The LP steam consumption also had come down to 9000 Kg/h.
With sincere efforts and study of operating data of plant. the energy has been brought down very
close to the guaranteed value of 660 Kcal/Nm3 of CO2. It is expected that once the suspected
leakage in LP stripper is stopped and the ejector performance is improved, the regeneration
energy may come down much below the guaranteed value.

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