Jones D.S.J., Pujado P.R. Handbook of Petroleum Processing
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442 CHAPTER 10
A typical heat flux over a kettle reboiler is 26,500 Btu/hr sqft therefore area for heat
transfer = 80 sqft.
The vapor generated by the reclaimer will be added to the vapor load at the bottom tray.
Likewise this amount as liquid will be added to the trays liquid load. The condenser
0.5 0.6 0.7 0.8 0.9 1.0 1.5 2.0 2.5 3.0 3.5 4.0
0
4.03.53.02.52.01.51.00.90.80.70.60.5
0
10.
20.
30.
40.
50.
60.
70.
80.
90.
91.
92.
93.
94.
95.
96.
97.
98.
98.5
99.
99.1
99.2
99.3
99.4
99.5
99.6
99.7
99.8
99.9
A or S
A = ABSORPTION FACTOR =
S = STRIPPING FACTOR =
10
20
30
40
50
60
70
80
90
91
92
93
94
95
96
97
98
99
99
98
97
96
95
94
93
92
91
90
L
YK
YK
L
PER CENT
ABSORPTION
RECOVERY
OR
PER CENT
STRIPPED
1 PLATE
PLATES
2
3
4
5
6
8
10
12
14
16
18
30
20
8
Figure 10.4. The Kremser equation and correlation.
REFINERY GAS TREATING PROCESSES 443
Table 10.10.
Component Moles/hr Ave K VK/l Strip out % Moles/ hr
H
2
S 270 44 4.7 100 270
H
2
O 8,655 0.72 0.08 < 0.5 8.0
HC 1.8 3.5 0.37 50 0.9
MEA 637.6 0.1 0.01 0 0
Total 8,564.3 278.9
Notes Use 3 trays From Figure 10.4
duty will also be increased to accommodate the reclaimer duty. Thus:
Total condenser duty = 8.125 + 2.130 = 10.255 mmBtu/hr
Vapor load to bottom tray = Stripout + reclaimed vapor
= 25,748 lbs/hr + 2851
= 28,599 lbs/hr
Moles/hr = 1,115.6 + 135.8
= 1,251.4
Stripper tower dimensions
As in the case of the contactor the cross sectional area of the stripper will be calculated
using the Brown and Souder method as follows:
Cross sectional area of the tower will be based on the loadings over the bottom tray
thus:
Total vapor to tray = 1,251.4 moles /hr
In actual cubic feet/sec =
1251.4 × 378 × 700 × 14.7
520 × 34 × 3,600
= 76.47 ACFS
ρ
v
= 7.85/76.47 = 0.104 lbs/cuft
Table 10.11.
Stream V or L
◦
F lbs/hr Btu/lb mmBtu/hr
IN
Feed L 249 3,930 230 0.935
Reclaimer By diff 2.130
Total in 3,930 3.065
OUT
Rec vapor V 260 2,851 1070 3.051
Sludge L 260 79 182 0.014
Total out 3,930 3.065
444 CHAPTER 10
Total liquid from the bottom tray = Stripout + Product + reclaimer liq
= 25,748 + 196,482 +2,851
= 225,081 lbs/hr
Lbs/gal of liquid at 240
◦
F = 7.87
Gals per hour = 28,600 and cubic ft/hr is 3,823
ρ
l
= 58.87 lbs/cuft.
Loading at flood = K
f
√
{ρ
v
× (ρ
l
− ρ
v
)}
From Figure 3.8.3 K
f
= 1,280 and at 60% = 768
Then loading = 768 × 2.47
= 1,899 lbs/hr sqft
let loading be 80% flood = 1,519
Cross sectional area =
28,599
1519
= 18.8 sqft and diameter
= 4.9 ft call it 5 ft or 19.6 sqft
Tower will contain 20 valve type trays made up of 18 stripping trays and 2 rectifying
trays. As in the case of the contactor trays they will be spaced at 30 inches. Trayed
section will therefore be 19 × 30 inches = 570 inches or 47 ft 6 inches from bottom
to top tray.
The bottom of the tower (from bottom tray to bottom tan will be sized to cater for a
2 min hold up of liquid to the HLL. Thus:
Total liquid from bottom reboiler = 196,482 + 2,851
= 199,333 lbs/hr
lbs/gal at 60
◦
F = 8.35 at 249 = 7.89 lbs/gal
Then gals/hr of total bottoms = 199,333/7.891
= 25,264 gals/hr
= 421 gpm
Or 56.3 cubic ft/min
Volume resident over three min = 168.9 say 169 cubic ft
Height of HLL = 8.6 allow 4 ft from HLL to bottom tray.
Stream
Comp
Moles/hr
Lbs/hr
H2S
H2O
HC
MEA
Total
GPM
(mmSCFD)
12 3
Amine
BEW
45 6
Sour
Recycle Gas
212.5
3094.4
3306.9
3092.6
3092.6
34722
Nil Nil
Nil
Nil Nil
454.6
8952
NilNil
(30) (28)
Sweet
Recycle Gas
(neg)
27462
Lean
Amine
57.5
637.5
8646.6
9341.6
196482
391 412
270 212.5
8655
1.8 1.8
8.4
29.5
5.5
419.6
637.5
9364.3
203988
Rich Amine Acid Gas
222.7
7506
(0.2)
Stripper
Reflux
15
8-C-1
Contactor
48''id × 696"T-T
20
8-C-1
1
312 PGL 105°F
2
1
100°F 320 TSG
LC
LC
105°F
Water
3
4
1-1
Surge Tank.
8-C-2
Stripper
60'' id × 768'' T-T
8–E-1
Lean/Rich Amine exchanger
13.44 mm Btu/hr
8-D-2
Reflux Drum
40'' id × 108'' T-T
8-E-4
Condenser
10.26 mmBtu/hr
216°F
27 psin
8-E-2
Lean Amine Trim Cooler
10.3 mm Btu/hr
.30°F
8-E-3
Stripper Reboiler
22.05 mm Btu/hr
8–E-S
Re-claimer
2.13 mm Btu/hr
Sludes
260 °F
175 °F
204 °F
248 °F
8-C-2
8
20
34 PSIA
8-T-I
Lc
Lc
Lc
Lc
b
5
To sulfur
Plant
Cease
8PI
100%
2 point
Figure 10.5. The preliminary process flow sheet for a MEA treating plant.
446 CHAPTER 10
Summary of Stripper height (all dimensions from bottom tan)
To bottom tray 12 ft 6 ins
Trayed section 47 ft 6 ins
Top tray to top tan 4 ft
Total height 64 ft
There will be a 15 ft skirt to give the total height above grade of 79 ft.
This completes the process design of the major items of equipment. These and the
smaller items of equipment are shown in the following preliminary process flow sheet
Figure 10.5.
Chapter 11
Upgrading the ‘Bottom of the Barrel’
D.S.J. Jones
The highest yield of straight run product as a percentage on crude is the residue
from the atmospheric distillation of the crude feed. In most Middle East crudes this
ranges from around 40 vol% for the lighter crude oils to 50% and higher in case of
the heavier oils. It can be, and often is, the major factor in crude oil slate selection
and in the operation of the refinery itself. In most cases a large portion of the residue
can be blended off to meet the fuel oil product in the refinery’s production plan. In
other cases, perhaps equally as common, a major portion of the atmospheric residue is
further distilled, under vacuum, to distillates which can be processed to meet gasoline
and middle distillate product slates or lube oil blending stocks. The residue from this
vacuum distillation, now considerably smaller in volume, is routed to fuel oil or to a
bitumen product pool.
There are cases however where the quantities of both atmospheric and vacuum residues
are high enough to limit the refinery’s throughput or limit its production of the more
valuable products and thus limit the refinery’s profitability. In such cases, the conver-
sion of these residues becomes attractive, and, in some cases absolutely imperative.
This latter case refers to those refineries that have no or a very small fuel oil market.
The upgrading of these residues is accomplished by the indirect processing of the
atmospheric residue, processing the distillates, by catalytic cracking or hydrocrack-
ing and then thermally cracking the vacuum residue. Most of the processes involved
with the conversion of the vacuum distillates are described in other chapters of this
Handbook dealing with hydrocracking and fluid catalytic cracking.
Direct upgrading of the atmospheric residue by the thermal cracking processes has
long been the ‘work horse’of the industry in processing the ‘Bottom of the Barrel’.
However, some catalytic processes have been developed which demonstrate a more
efficient and effective method than the thermal cracking routes. This chapter deals
with the direct processing of crude oil residuum, and is divided into the following
parts:
447
448 CHAPTER 11
r
Thermal cracking
r
‘Deep oil’fluid catalytic cracking
r
Residuum hydrocracking and desulfurization
The thermal cracking processes
Thermal cracking processes refer to those that convert the residuum feed (whether
atmospheric or vacuum residues) into higher grade products such as naphtha and
middle distillates, by heat at high temperature alone. That is, no catalyst or chemicals
are used in the conversion. The processes themselves are:
r
Visbreaking
r
Thermal cracking
r
Coking
Certain confusion exists in the definition of visbreaking and thermal cracking. Dif-
ferentiation is based on the type of feedstock, severity of cracking or the final result.
Strictly speaking, the term visbreaking should refer strictly to the viscosity reduction
of heavy stock as the process’s main objective.
Applications of the thermal cracker processes
In a simple refinery, without vacuum distillation facilities, the residue from the crude
oil atmospheric distillation typically boiling above 650
◦
For700
◦
F constitutes the
bulk of the heavy fuel oil produced. In cases where incremental production of light
and middle distillates at the expense of fuel oil is desired, one stage thermal cracking
of the residue is an easy and cost effective solution. The residue is cracked in a
specially designed heater, the effluent from the heater is quenched and routed to a
fractionator, sometimes with a pre-flash and the products of cracking such as light
gases, naphtha, gas oil, and residue are separated in the conventional manner. Some
20% of the residue feed can be converted into lighter products, mostly gas oil, by this
process. Figure 11.1 shows a typical one stage thermal cracker.
For increased gas oil production, a somewhat more complicated scheme can be ap-
plied. The residue feed in this case is first cracked in the heater and the effluent flashed,
the hot vapors from the flash drum are routed to a fractionating tower where a heavy
gas oil is recovered as the bottom product. This is, in turn, cracked in a second heater
and under more severe conditions to yield additional quantities of light distillate prod-
ucts and gas oil. Typically the first, residue, heater is operated at 15–20 psig and a coil
outlet temperature of 900
◦
F. The second, gas oil, heater operates at around 250–300
psig and a coil outlet temperature of 930
◦
F. This process is a two stage thermal cracker
and is shown as Figure 11.2.
UPGRADING RESIDUES 449
Figure 11.1. One stage thermal cracker.
The visbreaker process
The visbreaker process configuration is very similar to the single stage thermal cracker
as shown in Figure 11.1. More often than not though an additional piece of equipment
is added immediately after the heater. This is a simple soaking drum which prolongs
the time the heater effluent remains at the cracking temperature without being sub-
jected to further heat input and temperature. The objective here is to maintain good
fuel oil stability while still converting sufficient of the feed to gas oil and thus lowering
the residue viscosity to fuel oil specification. By providing a soaker drum suitable con-
version is obtained by residence time at moderate temperature and pressure conditions
Figure 11.2. Two stage thermal cracker.
450 CHAPTER 11
Figure 11.3. A typical visbreaker.
without impairing the resulting fuel oil stability. Visbreaker heater outlet temperature
can be as low as 830
◦
F to meet viscosity specification. This process configuration is
shown as Figure 11.3.
The coking processes
Coking is the most severe form of thermal cracking, in which all the residue feed
is converted into light ends, naphtha, middle distillate, and a solid material, coke.
No residual fuel oil is left. This process is particularly useful in handling very heavy
crudes such as Bachequero or the Canadian ‘Heavy Oil’crudes. A suitable market for
the coke must be found, and coke quality requirements may impose further processing
such as calcining with, of course, a resulting additional cost.
Two major types of coking processes are in operation today, these are the delayed
coking process and the fluid coking process. The intricacies, delicate operation and
cost of the conventional fluid process makes it an unlikely contender when simple
and cheap ways to convert residues into more valuable products are desired. A fairly
recent development of the fluid coking process, however, allows for the conversion of
the coke into low Btu gas. Only the delayed coking process and a proprietary process
for Fluid coking will be discussed here.
Delayed coking process
The delayed coking process is illustrated by Figure 11.4.
The fuel oil or heavy oil feed is routed to a cracking furnace similar to other thermal
cracking processes. The effluent from the furnace is sent to one of a set of several
UPGRADING RESIDUES 451
Figure 11.4. A delayed coker.
coking drums without quenching. The effluent is normally at a coil outlet temperature
of around 920
◦
F and at a pressure of 30–50 psig and the coking drum is filled in about
24 hr. The vapors leaving the top of the drum whilst filling are routed to the fractionator,
where they are fractionated to the distillate products. The very long residence period
of the effluent in the coking drums results in the complete destruction of the heavy
fractions, and a solid residue rich in carbon is left in the drum as coke.
When the drum is full of coke it is cooled and the top and bottom of the drum are
opened. Special high pressure water jets are used to cut and remove the coke through
the drum’s bottom opening.
The following table gives a rough comparison of yields from the various thermal
cracking processes. The two stage thermal cracker as shown in Figure 11.2 is consid-
ered in Table 11.1.
Table 11.1. Comparison of yields from thermal cracking processes,
wt% on feed
Thermal cracker Visbreaker Delayed coker
Feedstock Atmos residue Vacuum residue Atmos residue
Butane and lighter 4.0 2.5 7.5
C5–330
◦
F naphtha 7.0 4.5 15.5
330–660
◦
F gas oil 26.0 13.0 59.0
Cracked residue 63.0 80.0 18.0 As coke